Liquefied Natural Gas and Hydrocarbon Gas Processing

ABSTRACT

A process for the recovery of ethane, ethylene, propane, propylene, and heavier hydrocarbons from a liquefied natural gas (LNG) stream and a hydrocarbon gas stream is disclosed. The LNG feed stream is divided into two portions. The first portion is supplied to a fractionation column at a first upper mid-column feed point. The second portion is directed in heat exchange relation with a first portion of a warmer distillation stream rising from the fractionation stages of the column, whereby the LNG feed stream is partially heated and the distillation stream is totally condensed. The condensed distillation stream is divided into a “lean” LNG stream and a reflux stream, whereupon the reflux stream is supplied to the column at a top column feed position. The second portion of the LNG feed stream is heated further to partially or totally vaporize it and thereafter supplied to the column at a first lower mid-column feed position. The gas stream is divided into two portions. The second portion is expanded to the operating pressure of the column, then both portions are directed in heat exchange relation with the lean LNG stream and the second portion of the warmer distillation stream, whereby both portions of the gas stream are cooled, the lean LNG stream is vaporized, and the second portion of the distillation stream is heated. The first portion of the gas stream, which has been cooled to substantial condensation, is supplied to the column at a second upper mid-column feed point, and the second portion is supplied to the column at a second lower mid-column feed point. The quantities and temperatures of the feeds to the column are effective to maintain the column overhead temperature at a temperature whereby the major portion of the desired components is recovered in the bottom liquid product from the column.

This invention relates to a process for the separation of ethane andheavier hydrocarbons or propane and heavier hydrocarbons from liquefiednatural gas (hereinafter referred to as LNG) combined with theseparation of a gas containing hydrocarbons to provide a volatilemethane-rich gas stream and a less volatile natural gas liquids (NGL) orliquefied petroleum gas (LPG) stream. The applicants claim the benefitsunder Title 35, United States Code, Section 119(e) of prior U.S.Provisional Application No. 61/053,814 which was filed on May 16, 2008.

BACKGROUND OF THE INVENTION

As an alternative to transportation in pipelines, natural gas at remotelocations is sometimes liquefied and transported in special LNG tankersto appropriate LNG receiving and storage terminals. The LNG can then bere-vaporized and used as a gaseous fuel in the same fashion as naturalgas. Although LNG usually has a major proportion of methane, i.e.,methane comprises at least 50 mole percent of the LNG, it also containsrelatively lesser amounts of heavier hydrocarbons such as ethane,propane, butanes, and the like, as well as nitrogen. It is oftennecessary to separate some or all of the heavier hydrocarbons from themethane in the LNG so that the gaseous fuel resulting from vaporizingthe LNG conforms to pipeline specifications for heating value. Inaddition, it is often also desirable to separate the heavierhydrocarbons from the methane and ethane because these hydrocarbons havea higher value as liquid products (for use as petrochemical feedstocks,as an example) than their value as fuel.

Although there are many processes which may be used to separate ethaneand/or propane and heavier hydrocarbons from LNG, these processes oftenmust compromise between high recovery, low utility costs, and processsimplicity (and hence low capital investment). U.S. Pat. Nos. 2,952,984;3,837,172; 5,114,451; and 7,155,931 describe relevant LNG processescapable of ethane or propane recovery while producing the lean LNG as avapor stream that is thereafter compressed to delivery pressure to entera gas distribution network. However, lower utility costs may be possibleif the lean LNG is instead produced as a liquid stream that can bepumped (rather than compressed) to the delivery pressure of the gasdistribution network, with the lean LNG subsequently vaporized using alow level source of external heat or other means. U.S. Pat. Nos.6,604,380; 6,907,752; 6,941,771; 7,069,743; and 7,216,507 and co-pendingapplication Ser. Nos. 11/749,268 and 12/060,362 describe such processes.

Economics and logistics often dictate that LNG receiving terminals belocated close to the natural gas transmission lines that will transportthe re-vaporized LNG to consumers. In many cases, these areas also haveplants for processing natural gas produced in the region to recover theheavier hydrocarbons contained in the natural gas. Available processesfor separating these heavier hydrocarbons include those based uponcooling and refrigeration of gas, oil absorption, and refrigerated oilabsorption. Additionally, cryogenic processes have become popularbecause of the availability of economical equipment that produces powerwhile simultaneously expanding and extracting heat from the gas beingprocessed. Depending upon the pressure of the gas source, the richness(ethane, ethylene, and heavier hydrocarbons content) of the gas, and thedesired end products, each of these processes or a combination thereofmay be employed.

The cryogenic expansion process is now generally preferred for naturalgas liquids recovery because it provides maximum simplicity with ease ofstartup, operating flexibility, good efficiency, safety, and goodreliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904;4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039;4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545;5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507;5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880;6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; andco-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; and12/206,230 describe relevant processes (although the description of thepresent invention is based on different processing conditions than thosedescribed in the cited U.S. patents).

The present invention is generally concerned with the integratedrecovery of ethylene, ethane, propylene, propane, and heavierhydrocarbons from such LNG and gas streams. It uses a novel processarrangement to integrate the heating of the LNG stream and the coolingof the gas stream to eliminate the need for a separate vaporizer and theneed for external refrigeration, allowing high C₂ component recoverywhile keeping the processing equipment simple and the capital investmentlow. Further, the present invention offers a reduction in the utilities(power and heat) required to process the LNG and gas streams, resultingin lower operating costs than other processes, and also offeringsignificant reduction in capital investment.

Heretofore, assignee's U.S. Pat. No. 7,216,507 has been used to recoverC₂ components and heavier hydrocarbon components in plants processingLNG, while assignee's U.S. Pat. No. 5,568,737 has been used to recoverC₂ components and heavier hydrocarbon components in plants processingnatural gas. Surprisingly, applicants have found that by integratingcertain features of the assignee's U.S. Pat. No. 7,216,507 inventionwith certain features of the assignee's U.S. Pat. No. 5,568,737,extremely high C₂ component recovery levels can be accomplished usingless energy than that required by individual plants to process the LNGand natural gas separately.

A typical analysis of an LNG stream to be processed in accordance withthis invention would be, in approximate mole percent, 92.2% methane,6.0% ethane and other C₂ components, 1.1% propane and other C₃components, and traces of butanes plus, with the balance made up ofnitrogen. A typical analysis of a gas stream to be processed inaccordance with this invention would be, in approximate mole percent,80.1% methane, 9.5% ethane and other C₂ components, 5.6% propane andother C₃ components, 1.3% iso-butane, 1.1% normal butane, 0.8% pentanesplus, with the balance made up of nitrogen and carbon dioxide. Sulfurcontaining gases are also sometimes present.

For a better understanding of the present invention, reference is madeto the following examples and drawings. Referring to the drawings:

FIG. 1 is a flow diagram of a base case natural gas processing plantusing LNG to provide its refrigeration;

FIG. 2 is a flow diagram of base case LNG and natural gas processingplants in accordance with U.S. Pat. Nos. 7,216,507 and 5,568,737,respectively;

FIG. 3 is a flow diagram of an LNG and natural gas processing plant inaccordance with the present invention; and

FIGS. 4 through 8 are flow diagrams illustrating alternative means ofapplication of the present invention to LNG and natural gas streams.

FIGS. 1 and 2 are provided to quantify the advantages of the presentinvention.

In the following explanation of the above figures, tables are providedsummarizing flow rates calculated for representative process conditions.In the tables appearing herein, the values for flow rates (in moles perhour) have been rounded to the nearest whole number for convenience. Thetotal stream rates shown in the tables include all non-hydrocarboncomponents and hence are generally larger than the sum of the streamflow rates for the hydrocarbon components. Temperatures indicated areapproximate values rounded to the nearest degree. It should also benoted that the process design calculations performed for the purpose ofcomparing the processes depicted in the figures are based on theassumption of no heat leak from (or to) the surroundings to (or from)the process. The quality of commercially available insulating materialsmakes this a very reasonable assumption and one that is typically madeby those skilled in the art.

For convenience, process parameters are reported in both the traditionalBritish units and in the units of the Système International d'Unites(SI). The molar flow rates given in the tables may be interpreted aseither pound moles per hour or kilogram moles per hour. The energyconsumptions reported as horsepower (HP) and/or thousand British ThermalUnits per hour (MBTU/Hr) correspond to the stated molar flow rates inpound moles per hour. The energy consumptions reported as kilowatts (kW)correspond to the stated molar flow rates in kilogram moles per hour.

FIG. 1 is a flow diagram showing the design of a processing plant torecover C₂+ components from natural gas using an LNG stream to providerefrigeration. In the simulation of the FIG. 1 process, inlet gas entersthe plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31.If the inlet gas contains a concentration of sulfur compounds whichwould prevent the product streams from meeting specifications, thesulfur compounds are removed by appropriate pretreatment of the feed gas(not illustrated). In addition, the feed stream is usually dehydrated toprevent hydrate (ice) formation under cryogenic conditions. Soliddesiccant has typically been used for this purpose.

The inlet gas stream 31 is cooled in heat exchanger 12 by heat exchangewith a portion (stream 72 a) of partially warmed LNG at −174° F. [−114°C.] and cool distillation stream 38 a at −107° F. [−77° C.]. The cooledstream 31 a enters separator 13 at −79° F. [−62° C.] and 584 psia [4,027kPa(a)] where the vapor (stream 34) is separated from the condensedliquid (stream 35). Liquid stream 35 is flash expanded through anappropriate expansion device, such as expansion valve 17, to theoperating pressure (approximately 430 psia [2,965 kPa(a)]) offractionation tower 20. The expanded stream 35 a leaving expansion valve17 reaches a temperature of −93° F. [−70° C.] and is supplied tofractionation tower 20 at a first mid-column feed point.

The vapor from separator 13 (stream 34) enters a work expansion machine10 in which mechanical energy is extracted from this portion of the highpressure feed. The machine 10 expands the vapor substantiallyisentropically to slightly above the tower operating pressure, with thework expansion cooling the expanded stream 34 a to a temperature ofapproximately −101° F. [−74° C.]. The typical commercially availableexpanders are capable of recovering on the order of 80-88% of the worktheoretically available in an ideal isentropic expansion. The workrecovered is often used to drive a centrifugal compressor (such as item11) that can be used to re-compress the heated distillation stream(stream 38 b), for example. The expanded stream 34 a is further cooledto −124° F. [−87° C.] in heat exchanger 14 by heat exchange with colddistillation stream 38 at −143° F. [−97° C.], whereupon the partiallycondensed expanded stream 34 b is thereafter supplied to fractionationtower 20 at a second mid-column feed point.

The demethanizer in tower 20 is a conventional distillation columncontaining a plurality of vertically spaced trays, one or more packedbeds, or some combination of trays and packing to provide the necessarycontact between the liquids falling downward and the vapors risingupward. The column also includes reboilers (such as reboiler 19) whichheat and vaporize a portion of the liquids flowing down the column toprovide the stripping vapors which flow up the column to strip theliquid product, stream 41, of methane and lighter components. Liquidproduct stream 41 exits the bottom of the tower at 99° F. [37° C.],based on a typical specification of a methane to ethane ratio of 0.020:1on a molar basis in the bottom product.

Overhead distillation stream 43 is withdrawn from the upper section offractionation tower 20 at −143° F. [−97° C.] and is divided into twoportions, streams 44 and 47. The first portion, stream 44, flows toreflux condenser 22 where it is cooled to −237° F. [−149° C.] andtotally condensed by heat exchange with a portion (stream 72) of thecold LNG (stream 71 a). Condensed stream 44 a enters reflux separator 23wherein the condensed liquid (stream 46) is separated from anyuncondensed vapor (stream 45). The liquid stream 46 from refluxseparator 23 is pumped by reflux pump 24 to a pressure slightly abovethe operating pressure of demethanizer 20 and stream 46 a is thensupplied as cold top column feed (reflux) to demethanizer 20. This coldliquid reflux absorbs and condenses the C₂ components and heavierhydrocarbon components from the vapors rising in the upper section ofdemethanizer 20.

The second portion (stream 47) of overhead vapor stream 43 combines withany uncondensed vapor (stream 45) from reflux separator 23 to form colddistillation stream 38 at −143° F. [−97° C.]. Distillation stream 38passes countercurrently to expanded stream 34 a in heat exchanger 14where it is heated to −107° F. [−77° C.] (stream 38 a), andcountercurrently to inlet gas in heat exchanger 12 where it is heated to47° F. [8° C.] (stream 38 b). The distillation stream is thenre-compressed in two stages. The first stage is compressor 11 driven byexpansion machine 10. The second stage is compressor 21 driven by asupplemental power source which compresses stream 38 c to sales linepressure (stream 38 d). After cooling to 126° F. [52° C.] in dischargecooler 22, stream 38 e combines with warm LNG stream 71 b to form theresidue gas product (stream 42). Residue gas stream 42 flows to thesales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet linerequirements.

The LNG (stream 71) from LNG tank 50 enters pump 51 at −251° F. [−157°C.]. Pump 51 elevates the pressure of the LNG sufficiently so that itcan flow through heat exchangers and thence to the sales gas pipeline.Stream 71 a exits the pump 51 at −242° F. [−152° C.] and 1364 psia[9,401 kPa(a)] and is divided into two portions, streams 72 and 73. Thefirst portion, stream 72, is heated as described previously to −174° F.[−114° C.] in reflux condenser 22 as it provides cooling to the portion(stream 44) of overhead vapor stream 43 from fractionation tower 20, andto 43° F. [6° C.] in heat exchanger 12 as it provides cooling to theinlet gas. The second portion, stream 73, is heated to 35° F. [2° C.] inheat exchanger 53 using low level utility heat. The heated streams 72 band 73 a recombine to form warm LNG stream 71 b at 40° F. [4° C.], whichthereafter combines with distillation stream 38 e to form residue gasstream 42 as described previously.

A summary of stream flow rates and energy consumption for the processillustrated in FIG. 1 is set forth in the following table:

TABLE I (FIG. 1) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] StreamMethane Ethane Propane Butanes+ Total 31 42,545 5,048 2,972 1,658 53,14534 33,481 1,606 279 39 36,221 35 9,064 3,442 2,693 1,619 16,924 4350,499 25 0 0 51,534 44 8,055 4 0 0 8,221 45 0 0 0 0 0 46 8,055 4 0 08,221 47 42,444 21 0 0 43,313 38 42,444 21 0 0 43,313 71 40,293 2,642491 3 43,689 72 27,601 1,810 336 2 29,927 73 12,692 832 155 1 13,762 4282,737 2,663 491 3 87,002 41 101 5,027 2,972 1,658 9,832 Recoveries*Ethane 65.37% Propane 85.83% Butanes+ 99.83% Power LNG Feed Pump 3,561HP [ 5,854 kW] Reflux Pump 23 HP [ 38 kW] Residue Gas Compressor 24,612HP [ 40,462 kW] Totals 28,196 HP [ 46,354 kW] Low Level Utility Heat LNGHeater 68,990 MBTU/Hr [ 44,564 kW] High Level Utility Heat DemethanizerReboiler 80,020 MBTU/Hr [ 51,689 kW] Specific Power HP-Hr/Lb. Mole 2.868[kW-Hr/kg mole] [ 4.715 ] *(Based on un-rounded flow rates)

The recoveries reported in Table I are computed relative to the totalquantities of ethane, propane, and butanes+ contained in the gas streambeing processed in the plant and in the LNG stream. Although therecoveries are quite high relative to the heavier hydrocarbons containedin the gas being processed (99.58%, 100.00%, and 100.00%, respectively,for ethane, propane, and butanes+), none of the heavier hydrocarbonscontained in the LNG stream are captured in the FIG. 1 process. In fact,depending on the composition of LNG stream 71, the residue gas stream 42produced by the FIG. 1 process may not meet all pipeline specifications.The specific power reported in Table I is the power consumed per unit ofliquid product recovered, and is an indicator of the overall processefficiency.

FIG. 2 is a flow diagram showing processes to recover C₂+ componentsfrom LNG and natural gas in accordance with U.S. Pat. Nos. 7,216,507 and5,568,737, respectively, with the processed LNG stream used to providerefrigeration for the natural gas plant. The processes of FIG. 2 havebeen applied to the same LNG stream and inlet gas stream compositionsand conditions as described previously for FIG. 1.

In the simulation of the FIG. 2 process, the LNG to be processed (stream71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51elevates the pressure of the LNG sufficiently so that it can flowthrough heat exchangers and thence to expansion machine 55. Stream 71 aexits the pump at −242° F. [−152° C.] and 1364 psia [9,401 kPa(a)] andis split into two portions, streams 75 and 76. The first portion, stream75, is expanded to the operating pressure (approximately 415 psia [2,859kPa(a)]) of fractionation column 62 by expansion valve 58. The expandedstream 75 a leaves expansion valve 58 at −238° F. [−150° C.] and isthereafter supplied to tower 62 at an upper mid-column feed point.

The second portion, stream 76, is heated to −79° F. [−62° C.] in heatexchanger 52 by cooling compressed overhead distillation stream 79 a at−70° F. [−57° C.] and reflux stream 82 at −128° F. [−89° C.]. Thepartially heated stream 76 a is further heated and vaporized in heatexchanger 53 using low level utility heat. The heated stream 76 b at −5°F. [−20° C.] and 1334 psia [9,195 kPa(a)] enters work expansion machine55 in which mechanical energy is extracted from this portion of the highpressure feed. The machine 55 expands the vapor substantiallyisentropically to the tower operating pressure, with the work expansioncooling the expanded stream 76 c to a temperature of approximately −107°F. [−77° C.] before it is supplied as feed to fractionation column 62 ata lower mid-column feed point.

The demethanizer in fractionation column 62 is a conventionaldistillation column containing a plurality of vertically spaced trays,one or more packed beds, or some combination of trays and packingconsisting of two sections. The upper absorbing (rectification) sectioncontains the trays and/or packing to provide the necessary contactbetween the vapors rising upward and cold liquid falling downward tocondense and absorb the ethane and heavier components; the lowerstripping (demethanizing) section contains the trays and/or packing toprovide the necessary contact between the liquids falling downward andthe vapors rising upward. The demethanizing section also includes one ormore reboilers (such as side reboiler 60 using low level utility heat,and reboiler 61 using high level utility heat) which heat and vaporize aportion of the liquids flowing down the column to provide the strippingvapors which flow up the column. The column liquid stream 80 exits thebottom of the tower at 54° F. [12° C.], based on a typical specificationof a methane to ethane ratio of 0.020:1 on a molar basis in the bottomproduct.

Overhead distillation stream 79 is withdrawn from the upper section offractionation tower 62 at −144° F. [−98° C.] and flows to compressor 56driven by expansion machine 55, where it is compressed to 807 psia[5,567 kPa(a)] (stream 79 a). At this pressure, the stream is totallycondensed as it is cooled to −128° F. [−89° C.] in heat exchanger 52 asdescribed previously. The condensed liquid (stream 79 b) is then dividedinto two portions, streams 83 and 82. The first portion (stream 83) isthe methane-rich lean LNG stream, which is pumped by pump 63 to 1270psia [8,756 kPa(a)] for subsequent vaporization in heat exchanger 12,heating stream 83 a to 40° F. [4° C.] as described in paragraph [0032]below to produce warm lean LNG stream 83 b.

The remaining portion of condensed liquid stream 79 b, reflux stream 82,flows to heat exchanger 52 where it is subcooled to −237° F. [−149° C.]by heat exchange with a portion of the cold LNG (stream 76) as describedpreviously. The subcooled stream 82 a is then expanded to the operatingpressure of demethanizer 62 by expansion valve 57. The expanded stream82 b at −236° F. [−149° C.] is then supplied as cold top column feed(reflux) to demethanizer 62. This cold liquid reflux absorbs andcondenses the C₂ components and heavier hydrocarbon components from thevapors rising in the upper rectification section of demethanizer 62.

In the simulation of the FIG. 2 process, inlet gas enters the plant at126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feedstream 31 is cooled in heat exchanger 12 by heat exchange with cold leanLNG (stream 83 a) at −116° F. [−82° C.], cool distillation stream 38 aat −96° F. [−71° C.], and demethanizer liquids (stream 39) at −3° F.[−20° C.]. The cooled stream 31 a enters separator 13 at −67° F. [−55°C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 33) is separatedfrom the condensed liquid (stream 35). Liquid stream 35 is flashexpanded through an appropriate expansion device, such as expansionvalve 17, to the operating pressure (approximately 375 psia [2,583kPa(a)]) of fractionation tower 20. The expanded stream 35 a leavingexpansion valve 17 reaches a temperature of −86° F. [−65° C.] and issupplied to fractionation tower 20 at a first lower mid-column feedpoint.

Vapor stream 33 from separator 13 is divided into two streams, 32 and34. Stream 32, containing about 22% of the total vapor, passes throughheat exchanger 14 in heat exchange relation with cold distillationstream 38 at −150° F. [−101° C.] where it is cooled to substantialcondensation. The resulting substantially condensed stream 32 a at −144°F. [−98° C.] is then flash expanded through an appropriate expansiondevice, such as expansion valve 16, to the operating pressure offractionation tower 20, cooling stream 32 b to −148° F. [−100° C.]before it is supplied to fractionation tower 20 at an upper mid-columnfeed point.

The remaining 78% of the vapor from separator 13 (stream 34) enters awork expansion machine 10 in which mechanical energy is extracted fromthis portion of the high pressure feed. The machine 10 expands the vaporsubstantially isentropically to the tower operating pressure, with thework expansion cooling the expanded stream 34 a to a temperature ofapproximately −100° F. [−73° C.]. The partially condensed expandedstream 34 a is thereafter supplied as feed to fractionation tower 20 ata second lower mid-column feed point.

The demethanizer in fractionation column 20 is a conventionaldistillation column containing a plurality of vertically spaced trays,one or more packed beds, or some combination of trays and packingconsisting of two sections. The upper absorbing (rectification) sectioncontains the trays and/or packing to provide the necessary contactbetween the vapors rising upward and cold liquid falling downward tocondense and absorb the ethane and heavier components; the lowerstripping (demethanizing) section contains the trays and/or packing toprovide the necessary contact between the liquids falling downward andthe vapors rising upward. The demethanizing section also includes one ormore reboilers (such as the side reboiler in heat exchanger 12 describedpreviously, and reboiler 19 using high level utility heat) which heatand vaporize a portion of the liquids flowing down the column to providethe stripping vapors which flow up the column. The column liquid stream40 exits the bottom of the tower at 85° F. [30° C.], based on a typicalspecification of a methane to ethane ratio of 0.020:1 on a molar basisin the bottom product, and combines with stream 80 to form the liquidproduct (stream 41).

Overhead distillation stream 38 is withdrawn from the upper section offractionation tower 20 at −150° F. [−101° C.]. It passescountercurrently to vapor stream 32 and recycle stream 36 a in heatexchanger 14 where it is heated to −96° F. [−71° C.] (stream 38 a), andcountercurrently to inlet gas stream 31 and recycle stream 36 in heatexchanger 12 where it is heated to 6° F. [−15° C.] (stream 38 b). Thedistillation stream is then re-compressed in two stages. The first stageis compressor 11 driven by expansion machine 10. The second stage iscompressor 21 driven by a supplemental power source which compressesstream 38 c to sales line pressure (stream 38 d). After cooling to 126°F. [52° C.] in discharge cooler 22, stream 38 e is divided into twoportions, stream 37 and recycle stream 36. Stream 37 combines with warmlean LNG stream 83 b to form the residue gas product (stream 42).Residue gas stream 42 flows to the sales gas pipeline at 1262 psia[8,701 kPa(a)], sufficient to meet line requirements.

Recycle stream 36 flows to heat exchanger 12 and is cooled to −102° F.[−75° C.] by heat exchange with cool lean LNG (stream 83 a), cooldistillation stream 38 a, and demethanizer liquids (stream 39) asdescribed previously. Stream 36 a is further cooled to −144° F. [−98°C.] by heat exchange with cold distillation stream 38 in heat exchanger14 as described previously. The substantially condensed stream 36 b isthen expanded through an appropriate expansion device, such as expansionvalve 15, to the demethanizer operating pressure, resulting in coolingof the total stream to −152° F. [−102° C.]. The expanded stream 36 c isthen supplied to fractionation tower 20 as the top column feed. Thevapor portion of stream 36 c combines with the vapors rising from thetop fractionation stage of the column to form distillation stream 38,which is withdrawn from an upper region of the tower as described above.

A summary of stream flow rates and energy consumption for the processillustrated in FIG. 2 is set forth in the following table:

TABLE II (FIG. 2) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]Stream Methane Ethane Propane Butanes+ Total 31 42,545 5,048 2,972 1,65853,145 33 36,197 2,152 429 64 39,690 35 6,348 2,896 2,543 1,594 13,45532 8,027 477 95 14 8,801 34 28,170 1,675 334 50 30,889 38 52,982 30 0 054,112 36 10,537 6 0 0 10,762 37 42,445 24 0 0 43,350 40 100 5,024 2,9721,658 9,795 71 40,293 2,642 491 3 43,689 75 4,835 317 59 0 5,243 7635,458 2,325 432 3 38,446 79 45,588 16 0 0 45,898 82 5,348 2 0 0 5,38583 40,240 14 0 0 40,513 80 53 2,628 491 3 3,176 42 82,685 38 0 0 83,86341 153 7,652 3,463 1,661 12,971 Recoveries* Ethane 99.51% Propane100.00% Butanes+ 100.00% Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNGProduct Pump 1,746 HP [ 2,870 kW] Residue Gas Compressor 31,674 HP [52,072 kW] Totals 36,981 HP [ 60,796 kW] Low Level Utility Heat LiquidFeed Heater 66,200 MBTU/Hr [ 42,762 kW] Demethanizer Reboiler 60 23,350MBTU/Hr [ 15,083 kW] Totals 89,550 MBTU/Hr [ 57,845 kW] High LevelUtility Heat Demethanizer Reboiler 19 20,080 MBTU/Hr [ 12,971 kW]Demethanizer Reboiler 61 3,400 MBTU/Hr [ 2,196 kW] Totals 23,480 MBTU/Hr[ 15,167 kW] Specific Power HP-Hr/Lb. Mole 2.851 [kW-Hr/kg mole] [ 4.687] *(Based on un-rounded flow rates)

Comparison of the recovery levels displayed in Tables I and II showsthat the liquids recovery of the FIG. 2 processes is much higher thanthat of the FIG. 1 process due to the recovery of the heavierhydrocarbon liquids contained in the LNG stream in fractionation tower62. The ethane recovery improves from 65.37% to 99.51%, the propanerecovery improves from 85.83% to 100.00%, and the butanes+ recoveryimproves from 99.83% to 100.00%. In addition, the process efficiency ofthe FIG. 2 processes is improved by about 1% in terms of the specificpower relative to the FIG. 1 process.

DESCRIPTION OF THE INVENTION EXAMPLE 1

FIG. 3 illustrates a flow diagram of a process in accordance with thepresent invention. The LNG stream and inlet gas stream compositions andconditions considered in the process presented in FIG. 3 are the same asthose in the FIG. 1 and FIG. 2 processes. Accordingly, the FIG. 3process can be compared with the FIG. 1 and FIG. 2 processes toillustrate the advantages of the present invention.

In the simulation of the FIG. 3 process, the LNG to be processed (stream71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51elevates the pressure of the LNG sufficiently so that it can flowthrough heat exchangers and thence to separator 54. Stream 71 a exitsthe pump at −242° F. [−152° C.] and 1364 psia [9,401 kPa(a)] and issplit into two portions, streams 72 and 73. The first portion, stream72, becomes stream 75 and is expanded to the operating pressure(approximately 415 psia [2,859 kPa(a)]) of fractionation column 62 byexpansion valve 58. The expanded stream 75 a leaves expansion valve 58at −238° F. [−150C.] and is thereafter supplied to tower 62 at an uppermid-column feed point.

The second portion, stream 73, is heated prior to entering separator 54so that all or a portion of it is vaporized. In the example shown inFIG. 3, stream 73 is first heated to −77° F. [−61 ° C.] in heatexchanger 52 by cooling compressed overhead distillation stream 79 a at−70° F. [−57° C.] and reflux stream 81 at −16° F. [−82° C.]. Thepartially heated stream 73 a becomes stream 76 and is further heated inheat exchanger 53 using low level utility heat. (High level utilityheat, such as the heating medium used in tower reboiler 61, is normallymore expensive than low level utility heat, so lower operating cost isusually achieved when use of low level heat, such as sea water, ismaximized and the use of high level utility heat is minimized.) Notethat in all cases exchangers 52 and 53 are representative of either amultitude of individual heat exchangers or a single multi-pass heatexchanger, or any combination thereof. (The decision as to whether touse more than one heat exchanger for the indicated heating services willdepend on a number of factors including, but not limited to, inlet LNGflow rate, heat exchanger size, stream temperatures, etc.)

The heated stream 76 a enters separator 54 at −5° F. [−20° C.] and 1334psia [9,195 kPa(a)] where the vapor (stream 77) is separated from anyremaining liquid (stream 78). Vapor stream 77 enters a work expansionmachine 55 in which mechanical energy is extracted from this portion ofthe high pressure feed. The machine 55 expands the vapor substantiallyisentropically to the tower operating pressure, with the work expansioncooling the expanded stream 77 a to a temperature of approximately −107°F. [−77° C.]. The work recovered is often used to drive a centrifugalcompressor (such as item 56) that can be used to re-compress the columnoverhead vapor (stream 79), for example. The partially condensedexpanded stream 77 a is thereafter supplied as feed to fractionationcolumn 62 at a lower mid-column feed point. The separator liquid (stream78), if any, is expanded to the operating pressure of fractionationcolumn 62 by expansion valve 59 before expanded stream 78 a is suppliedto fractionation tower 62 at a second lower mid-column feed point.

The demethanizer in fractionation column 62 is a conventionaldistillation column containing a plurality of vertically spaced trays,one or more packed beds, or some combination of trays and packing. Thefractionation tower 62 may consist of two sections. The upper absorbing(rectification) section contains the trays and/or packing to provide thenecessary contact between the vapors rising upward and cold liquidfalling downward to condense and absorb the ethane and heaviercomponents; the lower stripping (demethanizing) section contains thetrays and/or packing to provide the necessary contact between theliquids falling downward and the vapors rising upward. The demethanizingsection also includes one or more reboilers (such as side reboiler 60using low level utility heat, and reboiler 61 using high level utilityheat) which heat and vaporize a portion of the liquids flowing down thecolumn to provide the stripping vapors which flow up the column. Thecolumn liquid stream 80 exits the bottom of the tower at 54° F. [12°C.], based on a typical specification of a methane to ethane ratio of0.020:1 on a molar basis in the bottom product.

Overhead distillation stream 79 is withdrawn from the upper section offractionation tower 62 at −144° F. [−98° C.] and flows to compressor 56driven by expansion machine 55, where it is compressed to 805 psia[5,554 kPa(a)] (stream 79 a). At this pressure, the stream is totallycondensed as it is cooled to −116° F. [−82° C.] in heat exchanger 52 asdescribed previously. The condensed liquid (stream 79 b) is then dividedinto two portions, streams 83 and 81. The first portion (stream 83) isthe methane-rich lean LNG stream, which is pumped by pump 63 to 1275psia [8,791 kPa(a)] for subsequent vaporization in heat exchangers 14and 12, heating stream 83 a to −94° F. [−70° C.] and 40° F. [4° C.],respectively, as described in paragraphs [0047] and [0049] below toproduce warm lean LNG stream 83 c.

The remaining portion of condensed liquid stream 79 b, stream 81, flowsto heat exchanger 52 where it is subcooled to −237° F. [−149° C.] byheat exchange with a portion of the cold LNG (stream 73) as describedpreviously. The subcooled stream 81 a is then divided into two portions,streams 82 and 36. The first portion, reflux stream 82, is expanded tothe operating pressure of demethanizer 62 by expansion valve 57. Theexpanded stream 82 a at −236° F. [−149° C.] is then supplied as cold topcolumn feed (reflux) to demethanizer 62. This cold liquid reflux absorbsand condenses the C₂ components and heavier hydrocarbon components fromthe vapors rising in the upper rectification section of demethanizer 62.The disposition of the second portion, reflux stream 36 for demethanizer20, is described in paragraph [0050] below.

In the simulation of the FIG. 3 process, inlet gas enters the plant at126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feedstream 31 is divided into two portions, streams 32 and 33. The firstportion, stream 32, is cooled in heat exchanger 12 by heat exchange withcool lean LNG (stream 83 b) at −94° F. [−70° C.], cool distillationstream 38 a at −94° F. [−70° C.], and demethanizer liquids (stream 39)at −78° F. [−61° C.]. The partially cooled stream 32 a is further cooledfrom −89° F. [−67° C.] to −120° F. [−85° C.] in heat exchanger 14 byheat exchange with cold lean LNG (stream 83 a) at −97° F. [−72° C.] andcold distillation stream 38 at −144° F. [−98° C.]. Note that in allcases exchangers 12 and 14 are representative of either a multitude ofindividual heat exchangers or a single multi-pass heat exchanger, or anycombination thereof. (The decision as to whether to use more than oneheat exchanger for the indicated heating services will depend on anumber of factors including, but not limited to, inlet gas flow rate,heat exchanger size, stream temperatures, etc.) The substantiallycondensed stream 32 b is then flash expanded through an appropriateexpansion device, such as expansion valve 16, to the operating pressure(approximately 415 psia [2,861 kPa(a)]) of fractionation tower 20,cooling stream 32 c to −132° F. [−91° C.] before it is supplied tofractionation tower 20 at an upper mid-column feed point.

The second portion of feed stream 31, stream 33, enters a work expansionmachine 10 in which mechanical energy is extracted from this portion ofthe high pressure feed. The machine 10 expands the vapor substantiallyisentropically to a pressure slightly above the operating pressure offractionation tower 20, with the work expansion cooling the expandedstream 33 a to a temperature of approximately 92° F. [33° C.]. The workrecovered is often used to drive a centrifugal compressor (such as item11) that can be used to re-compress the heated distillation stream(stream 38 b), for example. The expanded stream 33 a is further cooledin heat exchanger 12 by heat exchange with cool lean LNG (stream 83 b),cool distillation stream 38 a, and demethanizer liquids (stream 39) asdescribed previously. The further cooled stream 33 b enters separator 13at −84° F. [−65° C.] and 423 psia [2,916 kPa(a)] where the vapor (stream34) is separated from the condensed liquid (stream 35).

Vapor stream 34 is cooled to −120° F. [−85° C.] in heat exchanger 14 byheat exchange with cold lean LNG (stream 83 a) and cold distillationstream 38 as described previously. The partially condensed stream 34 ais then supplied to fractionation tower 20 at a first lower mid-columnfeed point. Liquid stream 35 is flash expanded through an appropriateexpansion device, such as expansion valve 17, to the operating pressureof fractionation tower 20. The expanded stream 35 a leaving expansionvalve 17 reaches a temperature of −85° F. [−65° C.] and is supplied tofractionation tower 20 at a second lower mid-column feed point.

The second portion of subcooled stream 81 a, reflux stream 36, isexpanded to the operating pressure of demethanizer 20 by expansion valve15. The expanded stream 36 a at −236° F. [−149° C.] is then supplied ascold top column feed (reflux) to demethanizer 20. This cold liquidreflux absorbs and condenses the C₂ components and heavier hydrocarboncomponents from the vapors rising in upper rectification section 20 a ofdemethanizer 20.

The demethanizer in fractionation column 20 is a conventionaldistillation column containing a plurality of vertically spaced trays,one or more packed beds, or some combination of trays and packing. Thefractionation tower 20 may consist of two sections. The upper absorbing(rectification) section 20 a contains the trays and/or packing toprovide the necessary contact between the vapors rising upward and coldliquid falling downward to condense and absorb the ethane and heaviercomponents; the lower stripping (demethanizing) section 20 b containsthe trays and/or packing to provide the necessary contact between theliquids falling downward and the vapors rising upward. Demethanizingsection 20 b also includes one or more reboilers (such as the sidereboiler in heat exchanger 12 described previously, and reboiler 19using high level utility heat) which heat and vaporize a portion of theliquids flowing down the column to provide the stripping vapors whichflow up the column. The column liquid stream 40 exits the bottom of thetower at 95° F. [35° C.], based on a typical specification of a methaneto ethane ratio of 0.020:1 on a molar basis in the bottom product, andcombines with stream 80 to form the liquid product (stream 41).

Overhead distillation stream 38 is withdrawn from the upper section offractionation tower 20 at −144° F. [−98° C.]. It passes countercurrentlyto the first portion (stream 32 a) of inlet gas stream 31 and vaporstream 34 in heat exchanger 14 where it is heated to −94° F. [−70° C.](stream 38 a), and countercurrently to the first portion (stream 32) ofinlet gas stream 31 and expanded second portion (stream 33 a) in heatexchanger 12 where it is heated to 13° F. [−11° C.] (stream 38 b). Thedistillation stream is then re-compressed in two stages. The first stageis compressor 11 driven by expansion machine 10. The second stage iscompressor 21 driven by a supplemental power source which compressesstream 38 c to sales gas line pressure (stream 38 d). After cooling to126° F. [52° C.] in discharge cooler 22, stream 38 e combines with warmlean LNG stream 83 c to form the residue gas product (stream 42).Residue gas stream 42 flows to the sales gas pipeline at 1262 psia[8,701 kPa(a)], sufficient to meet line requirements.

A summary of stream flow rates and energy consumption for the processillustrated in FIG. 3 is set forth in the following table:

TABLE III (FIG. 3) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]Stream Methane Ethane Propane Butanes+ Total 31 42,545 5,048 2,972 1,65853,145 32 5,531 656 386 215 6,909 33 37,014 4,392 2,586 1,443 46,236 3432,432 1,703 255 29 35,166 35 4,582 2,689 2,331 1,414 11,070 36 7,720 20 0 7,773 38 50,165 24 0 0 51,078 40 100 5,026 2,972 1,658 9,840 7140,293 2,642 491 3 43,689 72/75 4,916 322 60 0 5,330 73/76 35,377 2,320431 3 38,359 77 35,377 2,320 431 3 38,359 78 0 0 0 0 0 79 45,682 14 0 045,990 81 13,162 4 0 0 13,251 83 32,520 10 0 0 32,739 82 5,442 2 0 05,478 80 53 2,630 491 3 3,177 42 82,685 34 0 0 83,817 41 153 7,656 3,4631,661 13,017 Recoveries* Ethane 99.55% Propane 100.00% Butanes+ 100.00%Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump 1,740 HP [2,861 kW] Residue Gas Compressor 24,852 HP [ 40,856 kW] Totals 30,153 HP[ 49,571 kW] Low Level Utility Heat Liquid Feed Heater 65,000 MBTU/Hr [41,987 kW] Demethanizer Reboiler 60 19,000 MBTU/Hr [ 12,273 kW] Totals84,000 MBTU/Hr [ 54,260 kW] High Level Utility Heat DemethanizerReboiler 19 41,460 MBTU/Hr [ 26,781 kW] Demethanizer Reboiler 61 8,400MBTU/Hr [ 5,426 kW] Totals 49,860 MBTU/Hr [ 32,207 kW] Specific PowerHP-Hr/Lb. Mole 2.316 [kW-Hr/kg mole] [ 3.808 ] *(Based on un-roundedflow rates)

The improvement offered by the FIG. 3 embodiment of the presentinvention is astonishing compared to the FIG. 1 and FIG. 2 processes.Comparing the recovery levels displayed in Table III above for the FIG.3 embodiment with those in Table I for the FIG. 1 process shows that theFIG. 3 embodiment of the present invention improves the ethane recoveryfrom 65.37% to 99.55%, the propane recovery from 85.83% to 100.00%, andthe butanes+ recovery from 99.83% to 100.00%. Further, comparing theutilities consumptions in Table III with those in Table I shows thatalthough the power required for the FIG. 3 embodiment of the presentinvention is approximately 7% higher than the FIG. 1 process, theprocess efficiency of the FIG. 3 embodiment of the present invention issignificantly better than that of the FIG. 1 process. The gain inprocess efficiency is clearly seen in the drop in the specific power,from 2.868 HP-Hr/Lb. Mole [4.715 kW-Hr/kg mole] for the FIG. 1 processto 2.316 HP-Hr/Lb. Mole [3.808 kW-Hr/kg mole] for the FIG. 3 embodimentof the present invention, an increase of more than 19% in the productionefficiency.

Comparing the recovery levels displayed in Table III for the FIG. 3embodiment with those in Table II for the FIG. 2 processes shows thatthe liquids recovery levels are essentially the same. However, comparingthe utilities consumptions in Table III with those in Table II showsthat the power required for the FIG. 3 embodiment of the presentinvention is about 18% lower than the FIG. 2 processes. This results inreducing the specific power from 2.851 HP-Hr/Lb. Mole [4.687 kW-Hr/kgmole] for the FIG. 2 processes to 2.316 HP-Hr/Lb. Mole [3.808 kW-H/kgmole] for the FIG. 3 embodiment of the present invention, an improvementof nearly 19% in the production efficiency.

There are six primary factors that account for the improved efficiencyof the present invention. First, compared to many prior art processes,the present invention does not depend on the LNG feed itself to directlyserve as the reflux for fractionation column 62. Rather, therefrigeration inherent in the cold LNG is used in heat exchanger 52 togenerate a liquid reflux stream (stream 82) that contains very little ofthe C₂ components and heavier hydrocarbon components that are to berecovered, resulting in efficient rectification in the absorbing sectionof fractionation tower 62 and avoiding the equilibrium limitations ofsuch prior art processes. Second, splitting the LNG feed into twoportions before feeding fractionation column 62 allows more efficientuse of low level utility heat, thereby reducing the amount of high levelutility heat consumed by reboiler 61. The cold portion of the LNG feed(stream 75 a) serves as a supplemental reflux stream for fractionationtower 62, providing partial rectification of the vapors in the expandedvapor and liquid streams (streams 77 a and 78 a, respectively) so thatheating and at least partially vaporizing the other portion (stream 73)of the LNG feed does not unduly increase the condensing load in heatexchanger 52. Third, using a portion of the cold LNG feed (stream 75 a)as a supplemental reflux stream allows using less top reflux (stream 82a) for fractionation tower 62. The lower top reflux flow, plus thegreater degree of heating using low level utility heat in heat exchanger53, results in less total liquid feeding fractionation column 62,reducing the duty required in reboiler 61 and minimizing the amount ofhigh level utility heat needed to meet the specification for the bottomliquid product from demethanizer 62.

Fourth, using the cold lean LNG stream 83 a to provide “free”refrigeration to the gas streams in heat exchangers 12 and 14 eliminatesthe need for a separate vaporization means (such as heat exchanger 53 inthe FIG. 1 process) to re-vaporize the LNG prior to delivery to thesales gas pipeline. Fifth, cooling a portion (stream 32) of inlet gasstream 31 to substantial condensation prior to expansion to theoperating pressure of demethanizer 20 allows the expanded substantiallycondensed stream 32 c to serve as a supplemental reflux stream forfractionation tower 20, providing partial rectification of the vapors inthe partially condensed vapor and expanded liquid streams (streams 34 aand 35 a, respectively) so that less top reflux (stream 36 a) is neededfor fractionation tower 20. Sixth, integrating the LNG plant with thegas plant allows using a portion (stream 36) of the lean LNG as refluxfor demethanizer 20. The resulting stream 36 a is very cold and containsvery little of the C₂ components and heavier hydrocarbon components thatare to be recovered, resulting in very efficient rectification inabsorbing section 20 a and further minimizing the quantity of refluxrequired for demethanizer 20.

EXAMPLE 2

An alternative method of processing natural gas is shown in anotherembodiment of the present invention as illustrated in FIG. 4. The LNGstream and inlet gas stream compositions and conditions considered inthe process presented in FIG. 4 are the same as those in FIGS. 1 through3. Accordingly, the FIG. 4 process can be compared with the FIGS. 1 and2 processes to illustrate the advantages of the present invention, andcan likewise be compared to the embodiment displayed in FIG. 3.

In the simulation of the FIG. 4 process, the LNG to be processed (stream71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51elevates the pressure of the LNG sufficiently so that it can flowthrough heat exchangers and thence to separator 54. Stream 71 a exitsthe pump at −242° F. [−152° C.] and 1364 psia [9,401 kPa(a)] and issplit into two portions, streams 72 and 73. The first portion, stream72, becomes stream 75 and is expanded to the operating pressure(approximately 415 psia [2,859 kPa(a)]) of fractionation column 62 byexpansion valve 58. The expanded stream 75 a leaves expansion valve 58at −238° F. [−150C.] and is thereafter supplied to tower 62 at an uppermid-column feed point.

The second portion, stream 73, is heated prior to entering separator 54so that all or a portion of it is vaporized. In the example shown inFIG. 4, stream 73 is first heated to −77° F. [−61 ° C.] in heatexchanger 52 by cooling compressed overhead distillation stream 79 a at−70° F. [−57° C.] and reflux stream 81 at −15° F. [−82° C]. Thepartially heated stream 73 a becomes stream 76 and is further heated inheat exchanger 53 using low level utility heat. The heated stream 76 aenters separator 54 at −5° F. [−20° C.] and 1334 psia [9,195 kPa(a)]where the vapor (stream 77) is separated from any remaining liquid(stream 78). Vapor stream 77 enters a work expansion machine 55 in whichmechanical energy is extracted from this portion of the high pressurefeed. The machine 55 expands the vapor substantially isentropically tothe tower operating pressure, with the work expansion cooling theexpanded stream 77 a to a temperature of approximately −107° F. [−77°C.]. The partially condensed expanded stream 77 a is thereafter suppliedas feed to fractionation column 62 at a lower mid-column feed point. Theseparator liquid (stream 78), if any, is expanded to the operatingpressure of fractionation column 62 by expansion valve 59 beforeexpanded stream 78 a is supplied to fractionation tower 62 at a secondlower mid-column feed point.

The column liquid stream 80 exits the bottom of the tower at 54° F. [12°C.], based on a typical specification of a methane to ethane ratio of0.020:1 on a molar basis in the bottom product. Overhead distillationstream 79 is withdrawn from the upper section of fractionation tower 62at −144° F. [−98° C.] and flows to compressor 56 driven by expansionmachine 55, where it is compressed to 805 psia [5,554 kPa(a)] (stream 79a). At this pressure, the stream is totally condensed as it is cooled to−115° F. [−82° C.] in heat exchanger 52 as described previously. Thecondensed liquid (stream 79 b) is then divided into two portions,streams 83 and 81. The first portion (stream 83) is the methane-richlean LNG stream, which is pumped by pump 63 to 1270 psia [8,756 kPa(a)]for subsequent vaporization in heat exchanger 12, heating stream 83 a to40° F. [4° C.] as described in paragraph [0063] below to produce warmlean LNG stream 83 b.

The remaining portion of condensed liquid stream 79 b, stream 81, flowsto heat exchanger 52 where it is subcooled to −237° F. [−149° C.] byheat exchange with a portion of the cold LNG (stream 73) as describedpreviously. The subcooled stream 81 a is then divided into two portions,streams 82 and 36. The first portion, reflux stream 82, is expanded tothe operating pressure of demethanizer 62 by expansion valve 57. Theexpanded stream 82 a at −236° F. [−149° C.] is then supplied as cold topcolumn feed (reflux) to demethanizer 62. This cold liquid reflux absorbsand condenses the C₂ components and heavier hydrocarbon components fromthe vapors rising in the upper rectification section of demethanizer 62.The disposition of the second portion, reflux stream 36 for demethanizer20, is described in paragraph [0066] below.

In the simulation of the FIG. 4 process, inlet gas enters the plant at126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feedstream 31 is divided into two portions, streams 32 and 33. The firstportion, stream 32, is cooled in heat exchanger 12 by heat exchange withcold lean LNG (stream 83 a) at −96° F. [−71° C.], cool compresseddistillation stream 38 b at −109° F. [−78° C.], and demethanizer liquids(stream 39) at −63° F. [−53° C.]. The partially cooled stream 32 a isfurther cooled from −96° F. [−71° C.] to −121° F. [−85° C.] in heatexchanger 14 by heat exchange with cold compressed distillation stream38 a at −128° F. [−89° C.]. The substantially condensed stream 32 b isthen flash expanded through an appropriate expansion device, such asexpansion valve 16, to the operating pressure (approximately 443 psia[3,052 kPa(a)]) of fractionation tower 20, cooling stream 32 c to −129°F. [−90° C.] before it is supplied to fractionation tower 20 at an uppermid-column feed point.

The second portion of feed stream 31, stream 33, is cooled in heatexchanger 12 by heat exchange with cold lean LNG (stream 83 a), coolcompressed distillation stream 38 b, and demethanizer liquids (stream39) as described previously. The cooled stream 33 a enters separator 13at −86° F. [−65° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream34) is separated from the condensed liquid (stream 35). Liquid stream 35is flash expanded through an appropriate expansion device, such asexpansion valve 17, to the operating pressure of fractionation tower 20.The expanded stream 35 a leaving expansion valve 17 reaches atemperature of −100° F. [−73° C.] and is supplied to fractionation tower20 at a first lower mid-column feed point.

The vapor from separator 13 (stream 34) enters a work expansion machine10 in which mechanical energy is extracted from this portion of the highpressure feed. The machine 10 expands the vapor substantiallyisentropically to slightly above the tower operating pressure, with thework expansion cooling the expanded stream 34 a to a temperature ofapproximately −106° F. [−77° C.]. The expanded stream 34 a is furthercooled to −121° F. [−85° C.] in heat exchanger 14 by heat exchange withcold compressed distillation stream 38 a as described previously,whereupon the partially condensed expanded stream 34 b is thereaftersupplied to fractionation tower 20 at a second lower mid-column feedpoint.

The second portion of subcooled stream 81 a, reflux stream 36, isexpanded to the operating pressure of demethanizer 20 by expansion valve15. The expanded stream 36 a at −236° F. [−149° C.] is then supplied ascold top column feed (reflux) to demethanizer 20. This cold liquidreflux absorbs and condenses the C₂ components and heavier hydrocarboncomponents from the vapors rising in the upper rectification section ofdemethanizer 20.

The column liquid stream 40 exits the bottom of the tower at 102° F.[39° C.], based on a typical specification of a methane to ethane ratioof 0.020:1 on a molar basis in the bottom product, and combines withstream 80 to form the liquid product (stream 41). Overhead distillationstream 38 is withdrawn from the upper section of fractionation tower 20at −141° F. [−96° C.] and flows to compressor 11 driven by expansionmachine 10, where it is compressed to 501 psia [3,452 kPa(a)]. The coldcompressed distillation stream 38 a passes countercurrently to the firstportion (stream 32 a) of inlet gas stream 31 and expanded vapor stream34 a in heat exchanger 14 where it is heated to −109° F. [−78° C.](stream 38 b), and countercurrently to the first portion (stream 32) andsecond portion (stream 33) of inlet gas stream 31 in heat exchanger 12where it is heated to 31° F. [−1° C.] (stream 38 c). The heateddistillation stream then enters compressor 21 driven by a supplementalpower source which compresses stream 38 c to sales line pressure (stream38 d). After cooling to 126° F. [52° C.] in discharge cooler 22, stream38 e combines with warm lean LNG stream 83 b to form the residue gasproduct (stream 42). Residue gas stream 42 flows to the sales gaspipeline at 1262 psia [8,701 kPa(a)], sufficient to meet linerequirements.

A summary of stream flow rates and energy consumption for the processillustrated in FIG. 4 is set forth in the following table:

TABLE IV (FIG. 4) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]Stream Methane Ethane Propane Butanes+ Total 31 42,545 5,048 2,972 1,65853,145 32 3,404 404 238 133 4,251 33 39,141 4,644 2,734 1,525 48,894 3428,606 1,181 191 26 30,730 35 10,535 3,463 2,543 1,499 18,164 36 8,046 20 0 8,101 38 50,491 27 0 0 51,413 40 100 5,023 2,972 1,658 9,833 7140,293 2,642 491 3 43,689 72/75 4,916 322 60 0 5,330 73/76 35,377 2,320431 3 38,359 77 35,377 2,320 431 3 38,359 78 0 0 0 0 0 79 45,682 14 0 045,990 81 13,488 4 0 0 13,579 83 32,194 10 0 0 32,411 82 5,442 2 0 05,478 80 53 2,630 491 3 3,177 42 82,685 37 0 0 83,824 41 153 7,653 3,4631,661 13,010 Recoveries* Ethane 99.51% Propane 100.00% Butanes+ 100.00%Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump 1,727 HP [2,839 kW] Residue Gas Compressor 24,400 HP [ 40,113 kW] Totals 29,688 HP[ 48,806 kW] Low Level Utility Heat Liquid Feed Heater 65,000 MBTU/Hr [41,987 kW] Demethanizer Reboiler 60 19,000 MBTU/Hr [ 12,273 kW] Totals84,000 MBTU/Hr [ 54,260 kW] High Level Utility Heat DemethanizerReboiler 19 37,360 MBTU/Hr [ 24,133 kW] Demethanizer Reboiler 61 8,400MBTU/Hr [ 5,426 kW] Totals 45,760 MBTU/Hr [ 29,559 kW] Specific PowerHP-Hr/Lb. Mole 2.282 [kW-Hr/kg mole] [ 3.751 ] *(Based on un-roundedflow rates)

A comparison of Tables III and IV shows that the FIG. 4 embodiment ofthe present invention achieves essentially the same liquids recovery asthe FIG. 3 embodiment. However, the FIG. 4 embodiment uses less powerthan the FIG. 3 embodiment, improving the specific power by slightlymore than 1%. In addition, the high level utility heat required for theFIG. 4 embodiment of the present invention is about 8% less than that ofthe FIG. 3 embodiment.

EXAMPLE 3

Another alternative method of processing natural gas is shown in theembodiment of the present invention as illustrated in FIG. 5. The LNGstream and inlet gas stream compositions and conditions considered inthe process presented in FIG. 5 are the same as those in FIGS. 1 through4. Accordingly, the FIG. 5 process can be compared with the FIGS. 1 and2 processes to illustrate the advantages of the present invention, andcan likewise be compared to the embodiments displayed in FIGS. 3 and 4.

In the simulation of the FIG. 5 process, the LNG to be processed (stream71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51elevates the pressure of the LNG sufficiently so that it can flowthrough heat exchangers and thence to separator 54. Stream 71 a exitsthe pump at −242° F. [−152° C.] and 1364 psia [9,401 kPa(a)] and issplit into two portions, streams 72 and 73. The first portion, stream72, becomes stream 75 and is expanded to the operating pressure(approximately 415 psia [2,859 kPa(a)]) of fractionation column 62 byexpansion valve 58. The expanded stream 75 a leaves expansion valve 58at −238° F. [−150C] and is thereafter supplied to tower 62 at an uppermid-column feed point.

The second portion, stream 73, is heated prior to entering separator 54so that all or a portion of it is vaporized. In the example shown inFIG. 5, stream 73 is first heated to −77° F. [−61° C.] in heat exchanger52 by cooling compressed overhead distillation stream 79 a at −70° F.[−57° C.] and reflux stream 81 at −12° F. [−80° C]. The partially heatedstream 73 a becomes stream 76 and is further heated in heat exchanger 53using low level utility heat. The heated stream 76 a enters separator 54at −5° F. [−20° C.] and 1334 psia [9,195 kPa(a)] where the vapor (stream77) is separated from any remaining liquid (stream 78). Vapor stream 77enters a work expansion machine 55 in which mechanical energy isextracted from this portion of the high pressure feed. The machine 55expands the vapor substantially isentropically to the tower operatingpressure, with the work expansion cooling the expanded stream 77 a to atemperature of approximately −107° F. [−77° C.]. The partially condensedexpanded stream 77 a is thereafter supplied as feed to fractionationcolumn 62 at a lower mid-column feed point. The separator liquid (stream78), if any, is expanded to the operating pressure of fractionationcolumn 62 by expansion valve 59 before expanded stream 78 a is suppliedto fractionation tower 62 at a second lower mid-column feed point.

The column liquid stream 80 exits the bottom of the tower at 54° F. [12°C.], based on a typical specification of a methane to ethane ratio of0.020:1 on a molar basis in the bottom product. Overhead distillationstream 79 is withdrawn from the upper section of fractionation tower 62at −144° F. [−98° C.] and flows to compressor 56 driven by expansionmachine 55, where it is compressed to 805 psia [5,554 kPa(a)] (stream 79a). At this pressure, the stream is totally condensed as it is cooled to−112° F. [−80° C.] in heat exchanger 52 as described previously. Thecondensed liquid (stream 79 b) is then divided into two portions,streams 83 and 81. The first portion (stream 83) is the methane-richlean LNG stream, which is pumped by pump 63 to 1270 psia [8,756 kPa(a)]for subsequent vaporization in heat exchanger 12, heating stream 83 a to40° F. [4° C.] as described in paragraph [0075] below to produce warmlean LNG stream 83 b.

The remaining portion of condensed liquid stream 79 b, stream 81, flowsto heat exchanger 52 where it is subcooled to −237° F. [−149° C.] byheat exchange with a portion of the cold LNG (stream 73) as describedpreviously. The subcooled stream 81 a is then divided into two portions,streams 82 and 36. The first portion, reflux stream 82, is expanded tothe operating pressure of demethanizer 62 by expansion valve 57. Theexpanded stream 82 a at −236° F. [−149° C.] is then supplied as cold topcolumn feed (reflux) to demethanizer 62. This cold liquid reflux absorbsand condenses the C₂ components and heavier hydrocarbon components fromthe vapors rising in the upper rectification section of demethanizer 62.The disposition of the second portion, reflux stream 36 for demethanizer20, is described in paragraph [0078] below.

In the simulation of the FIG. 5 process, inlet gas enters the plant at126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feedstream 31 is divided into two portions, streams 32 and 33. The firstportion, stream 32, is cooled in heat exchanger 12 by heat exchange withcold lean LNG (stream 83 a) at −89° F. [−67° C.], cool compresseddistillation stream 38 b at −91° F. [−68° C.], and demethanizer liquids(stream 39) at −89° F. [−67° C.]. The partially cooled stream 32 a isfurther cooled from −86° F. [−65° C.] to −100° F. [−74° C.] in heatexchanger 14 by heat exchange with cold compressed distillation stream38 a at −112° F. [−80° C.]. The substantially condensed stream 32 b isthen flash expanded through an appropriate expansion device, such asexpansion valve 16, to the operating pressure (approximately 428 psia[2,949 kPa(a)]) of fractionation tower 20, cooling stream 32 c to −117°F. [−83° C.] before it is supplied to fractionation tower 20 at an uppermid-column feed point.

The second portion of feed stream 31, stream 33, enters a work expansionmachine 10 in which mechanical energy is extracted from this portion ofthe high pressure feed. The machine 10 expands the vapor substantiallyisentropically to a pressure slightly above the operating pressure offractionation tower 20, with the work expansion cooling the expandedstream 33 a to a temperature of approximately 95° F. [35° C.]. Theexpanded stream 33 a is further cooled in heat exchanger 12 by heatexchange with cold lean LNG (stream 83 a), cool compressed distillationstream 38 b, and demethanizer liquids (stream 39) as describedpreviously. The further cooled stream 33 b enters separator 13 at −85°F. [−65° C.] and 436 psia [3,004 kPa(a)] where the vapor (stream 34) isseparated from the condensed liquid (stream 35).

Vapor stream 34 is cooled to −100° F. [−74° C.] in heat exchanger 14 byheat exchange with cold compressed distillation stream 38 a as describedpreviously. The partially condensed stream 34 a is then supplied tofractionation tower 20 at a first lower mid-column feed point. Liquidstream 35 is flash expanded through an appropriate expansion device,such as expansion valve 17, to the operating pressure of fractionationtower 20. The expanded stream 35 a leaving expansion valve 17 reaches atemperature of −86° F. [−65° C.] and is supplied to fractionation tower20 at a second lower mid-column feed point.

The second portion of subcooled stream 81 a, reflux stream 36, isexpanded to the operating pressure of demethanizer 20 by expansion valve15. The expanded stream 36 a at −236° F. [−149° C.] is then supplied ascold top column feed (reflux) to demethanizer 20. This cold liquidreflux absorbs and condenses the C₂ components and heavier hydrocarboncomponents from the vapors rising in the upper rectification section ofdemethanizer 20.

The column liquid stream 40 exits the bottom of the tower at 98° F. [37°C.], based on a typical specification of a methane to ethane ratio of0.020:1 on a molar basis in the bottom product, and combines with stream80 to form the liquid product (stream 41). Overhead distillation stream38 is withdrawn from the upper section of fractionation tower 20 at−143° F. [−97° C.] and flows to compressor 11 driven by expansionmachine 10, where it is compressed to 573 psia [3,950 kPa(a)]. The coldcompressed distillation stream 38 a passes countercurrently to the firstportion (stream 32 a) of inlet gas stream 31 and vapor stream 34 in heatexchanger 14 where it is heated to −91° F. [−68° C.] (stream 38 b), andcountercurrently to the first portion (stream 32) and expanded secondportion (stream 33 a) of inlet gas stream 31 in heat exchanger 12 whereit is heated to 67° F. [19° C.] (stream 38 c). The heated distillationstream then enters compressor 21 driven by a supplemental power sourcewhich compresses stream 38 c to sales line pressure (stream 38 d). Aftercooling to 126° F. [52° C.] in discharge cooler 22, stream 38 e combineswith warm lean LNG stream 83 b to form the residue gas product (stream42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia[8,701 kPa(a)], sufficient to meet line requirements.

A summary of stream flow rates and energy consumption for the processillustrated in FIG. 5 is set forth in the following table:

TABLE V (FIG. 5) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr] StreamMethane Ethane Propane Butanes+ Total 31 42,545 5,048 2,972 1,658 53,14532 14,465 1,716 1,010 564 18,069 33 28,080 3,332 1,962 1,094 35,076 3424,317 1,236 184 21 26,322 35 3,763 2,096 1,778 1,073 8,754 36 10,372 30 0 10,442 38 52,817 30 0 0 53,749 40 100 5,021 2,972 1,658 9,838 7140,293 2,642 491 3 43,689 72/75 4,916 322 60 0 5,330 73/76 35,377 2,320431 3 38,359 77 35,377 2,320 431 3 38,359 78 0 0 0 0 0 79 45,682 14 0 045,990 81 15,814 5 0 0 15,920 83 29,868 9 0 0 30,070 82 5,442 2 0 05,478 80 53 2,630 491 3 3,177 42 82,685 39 0 0 83,819 41 153 7,651 3,4631,661 13,015 Recoveries* Ethane 99.48% Propane 100.00% Butanes+ 100.00%Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump 1,778 HP [2,923 kW] Residue Gas Compressor 23,201 HP [ 38,142 kW] Totals 28,540 HP[ 46,919 kW] Low Level Utility Heat Liquid Feed Heater 65,000 MBTU/Hr [41,987 kW] Demethanizer Reboiler 60 19,000 MBTU/Hr [ 12,273 kW] Totals84,000 MBTU/Hr [ 54,260 kW] High Level Utility Heat DemethanizerReboiler 19 53,370 MBTU/Hr [ 34,475 kW] Demethanizer Reboiler 61 8,400MBTU/Hr [ 5,426 kW] Totals 61,770 MBTU/Hr [ 39,901 kW] Specific PowerHP-Hr/Lb. Mole 2.193 [kW-Hr/kg mole] [ 3.605 ] *(Based on un-roundedflow rates)

A comparison of Tables III, IV, and V shows that the FIG. 5 embodimentof the present invention achieves essentially the same liquids recoveryas the FIG. 3 and FIG. 4 embodiments. The FIG. 5 embodiment uses lesspower than the FIG. 3 and FIG. 4 embodiments, improving the specificpower by over 5% relative to the FIG. 3 embodiment and nearly 4%relative to the FIG. 4 embodiment. However, the high level utility heatrequired for the FIG. 5 embodiment of the present invention is somewhathigher than that of the FIG. 3 and FIG. 4 embodiments (by 24% and 35%,respectively). The choice of which embodiment to use for a particularapplication will generally be dictated by the relative costs of powerand high level utility heat and the relative capital costs of pumps,heat exchangers, and compressors.

EXAMPLE 4

An alternative method of processing LNG and natural gas is shown in theembodiment of the present invention as illustrated in FIG. 6. The LNGstream and inlet gas stream compositions and conditions considered inthe process presented in FIG. 6 are the same as those in FIGS. 1 through5. Accordingly, the FIG. 5 process can be compared with the FIGS. 1 and2 processes to illustrate the advantages of the present invention, andcan likewise be compared to the embodiments displayed in FIGS. 3 through5.

In the simulation of the FIG. 6 process, the LNG to be processed (stream71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51elevates the pressure of the LNG sufficiently so that it can flowthrough heat exchangers and thence to separator 54. Stream 71 a exitsthe pump at −242° F. [−152° C.] and 1364 psia [9,401 kPa(a)] and issplit into two portions, streams 72 and 73. The first portion, stream72, becomes stream 75 and is expanded to the operating pressure(approximately 435 psia [2,997 kPa(a)]) of fractionation column 20 byexpansion valve 58. The expanded stream 75 a leaves expansion valve 58at −238° F. [−150° C.] and is thereafter supplied to tower 20 at a firstupper mid-column feed point.

The second portion, stream 73, is heated prior to entering separator 54so that all or a portion of it is vaporized. In the example shown inFIG. 6, stream 73 is first heated to −76° F. [−60° C.] in heat exchanger52 by cooling compressed overhead distillation stream 81 a at −65° F.[−54° C.] and reflux stream 82 at −117° F. [−82° C.], exchanger 14 asdescribed in paragraph [0085] below. The partially heated stream 73 bbecomes stream 76 and is further heated in heat exchanger 53 using lowlevel utility heat. The heated stream 76 a enters separator 54 at −5° F.[−20° C.] and 1334 psia [9,195 kPa(a)] where the vapor (stream 77) isseparated from any remaining liquid (stream 78). Vapor stream 77 entersa work expansion machine 55 in which mechanical energy is extracted fromthis portion of the high pressure feed. The machine 55 expands the vaporsubstantially isentropically to the tower operating pressure, with thework expansion cooling the expanded stream 77 a to a temperature ofapproximately −104° F. [−76° C.]. The partially condensed expandedstream 77 a is thereafter supplied as feed to fractionation column 20 ata first lower mid-column feed point. The separator liquid (stream 78),if any, is expanded to the operating pressure of fractionation column 20by expansion valve 59 before expanded stream 78 a is supplied tofractionation tower 20 at a second lower mid-column feed point.

In the simulation of the FIG. 6 process, inlet gas enters the plant at126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feedstream 31 is divided into two portions, streams 32 and 33. The firstportion, stream 32, is cooled in heat exchanger 12 by heat exchange withcold lean LNG (stream 83 a) at −103° F. [−75° C.], cool compresseddistillation stream 38 b at −92° F. [−69° C.], and demethanizer liquids(stream 39) at −78° F. [−61° C.]. The partially cooled stream 32 a isfurther cooled from −94° F. [−70° C.] to −101° F. [−74° C.] in heatexchanger 14 by heat exchange with the partially heated second portion(stream 73 a) of the LNG stream and with cold compressed distillationstream 38 a at −106° F. [−77° C.]. The substantially condensed stream 32b is then flash expanded through an appropriate expansion device, suchas expansion valve 16, to the operating pressure of fractionation tower20, cooling stream 32 c to −117° F. [−83° C.] before it is supplied tofractionation tower 20 at a second upper mid-column feed point.

The second portion of feed stream 31, stream 33, enters a work expansionmachine 10 in which mechanical energy is extracted from this portion ofthe high pressure feed. The machine 10 expands the vapor substantiallyisentropically to a pressure slightly above the operating pressure offractionation tower 20, with the work expansion cooling the expandedstream 33 a to a temperature of approximately 96° F. [36° C.]. Theexpanded stream 33 a is further cooled in heat exchanger 12 by heatexchange with cold lean LNG (stream 83 a), cool compressed distillationstream 38 b, and demethanizer liquids (stream 39) as describedpreviously. The further cooled stream 33 b enters separator 13 at −90°F. [−68° C.] and 443 psia [3,052 kPa(a)] where the vapor (stream 34) isseparated from the condensed liquid (stream 35).

Vapor stream 34 is cooled to −101° F. [−74° C.] in heat exchanger 14 byheat exchange with the partially heated second portion (stream 73 a) ofthe LNG stream and with cold compressed distillation stream 38 a asdescribed previously. The partially condensed stream 34 a is thensupplied to fractionation tower 20 at a third lower mid-column feedpoint. Liquid stream 35 is flash expanded through an appropriateexpansion device, such as expansion valve 17, to the operating pressureof fractionation tower 20. The expanded stream 35 a leaving expansionvalve 17 reaches a temperature of −90° F. [−68° C.] and is supplied tofractionation tower 20 at a fourth lower mid-column feed point.

The liquid product stream 41 exits the bottom of the tower at 89° F.[32° C.], based on a typical specification of a methane to ethane ratioof 0.020:1 on a molar basis in the bottom product. Overhead distillationstream 79 is withdrawn from the upper section of fractionation tower 20at −142° F. [−97° C.] and is divided into two portions, stream 81 andstream 38. The first portion (stream 81) flows to compressor 56 drivenby expansion machine 55, where it is compressed to 864 psia [5,955kPa(a)] (stream 81 a). At this pressure, the stream is totally condensedas it is cooled to −117° F. [−83° C.] in heat exchanger 52 as describedpreviously. The condensed liquid (stream 81 b) is then divided into twoportions, streams 83 and 82. The first portion (stream 83) is themethane-rich lean LNG stream, which is pumped by pump 63 to 1270 psia[8,756 kPa(a)] for subsequent vaporization in heat exchanger 12, heatingstream 83 a to 40° F. [4° C.] as described previously to produce warmlean LNG stream 83 b.

The remaining portion of stream 81 b (stream 82) flows to heat exchanger52 where it is subcooled to −237° F. [−149° C.] by heat exchange with aportion of the cold LNG (stream 73) as described previously. Thesubcooled stream 82 a is expanded to the operating pressure offractionation column 20 by expansion valve 57. The expanded stream 82 bat −236° F. [−149° C.] is then supplied as cold top column feed (reflux)to demethanizer 20. This cold liquid reflux absorbs and condenses the C₂components and heavier hydrocarbon components from the vapors rising inthe upper rectification section of demethanizer 20.

The second portion of distillation stream 79 (stream 38) flows tocompressor 11 driven by expansion machine 10, where it is compressed to604 psia [4,165 kPa(a)]. The cold compressed distillation stream 38 apasses countercurrently to the first portion (stream 32 a) of inlet gasstream 31 and vapor stream 34 in heat exchanger 14 where it is heated to−92° F. [−69° C.] (stream 38 b), and countercurrently to the firstportion (stream 32) and expanded second portion (stream 33 a) of inletgas stream 31 in heat exchanger 12 where it is heated to 48° F. [9° C.](stream 38 c). The heated distillation stream then enters compressor 21driven by a supplemental power source which compresses stream 38 c tosales line pressure (stream 38 d). After cooling to 126° F. [52° C.] indischarge cooler 22, stream 38 e combines with warm lean LNG stream 83 bto form the residue gas product (stream 42). Residue gas stream 42 flowsto the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient tomeet line requirements.

A summary of stream flow rates and energy consumption for the processillustrated in FIG. 6 is set forth in the following table:

TABLE VI (FIG. 6) Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]Stream Methane Ethane Propane Butanes+ Total 31 42,545 5,048 2,972 1,65853,145 32 7,871 934 550 307 9,832 33 34,674 4,114 2,422 1,351 43,313 3429,159 1,328 185 21 31,380 35 5,515 2,786 2,237 1,330 11,933 71 40,2932,642 491 3 43,689 72/75 5,037 330 61 0 5,461 73/76 35,256 2,312 430 338,228 77 35,256 2,312 430 3 38,228 78 0 0 0 0 0 79 97,329 46 0 0 98,69638 54,991 26 0 0 55,763 81 42,338 20 0 0 42,933 82 14,644 7 0 0 14,85083 27,694 13 0 0 28,083 42 82,685 39 0 0 83,846 41 153 7,651 3,463 1,66112,988 Recoveries* Ethane 99.48% Propane 100.00% Butanes+ 100.00% PowerLNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump 1,216 HP [ 1,999 kW]Residue Gas Compressor 21,186 HP [ 34,829 kW] Totals 25,963 HP [ 42,682kW] Low Level Utility Heat Liquid Feed Heater 70,000 MBTU/Hr [ 45,217kW] Demethanizer Reboiler 18 30,000 MBTU/Hr [ 19,378 kW] Totals 100,000MBTU/Hr [ 64,595 kW] High Level Utility Heat Demethanizer Reboiler 1939,180 MBTU/Hr [ 25,308 kW] Specific Power HP-Hr/Lb. Mole 1.999[kW-Hr/kg mole] [ 3.286 ] *(Based on un-rounded flow rates)

A comparison of Tables III, IV, V, and VI shows that the FIG. 6embodiment of the present invention achieves essentially the sameliquids recovery as the FIGS. 3, 4, and 5 embodiments. However, thereduction in the energy consumption of the FIG. 6 embodiment of thepresent invention relative to the embodiments in FIGS. 3 through 5 isunexpectedly large. The FIG. 6 embodiment uses less power than the FIGS.3, 4, and 5 embodiments, reducing the specific power by 14%, 12%, and9%, respectively. The high level utility heat required for the FIG. 6embodiment of the present invention is also lower than that of the FIGS.3, 4, and 5 embodiments (by 21%, 14%, and 37%, respectively). Theselarge gains in process efficiency are mainly due to the more optimaldistribution of the column feeds afforded by integrating the LNGprocessing and the natural gas processing into a single fractionationcolumn, demethanizer 20. For instance, the relative distribution of theinlet gas stream 31 between stream 32 (which forms the substantiallycondensed expanded stream 32 c) and stream 33 supplied to expansionmachine 10 can be optimized for power production, since stream 75 a fromLNG stream 71 provides part of the supplemental rectification for column20 that must be provided entirely by stream 32 c in the FIGS. 3 through5 embodiments.

The capital cost of the FIG. 6 embodiment of the present invention willgenerally be less than that of the FIGS. 3, 4, and 5 embodiments sinceit uses only one fractionation column, and due to the reduction in powerand high level utility heat consumption. The choice of which embodimentto use for a particular application will generally be dictated by therelative costs of power and high level utility heat and the relativecapital costs of columns, pumps, heat exchangers, and compressors.

Other Embodiments

Some circumstances may favor using cold distillation stream 38 in theFIG. 6 embodiment for heat exchange prior to compression as shown in theembodiment displayed in FIG. 7. In other instances, work expansion ofthe high pressure inlet gas may be more advantageous after cooling andseparation of any liquids, as shown in the embodiment displayed in FIG.8. The choices regarding the streams used for work expansion and wherebest to apply the power generated in compressing the process streamswill depend on such factors as inlet gas pressure and composition, andmust be determined for each application.

When the inlet gas is leaner, separator 13 in FIGS. 3 through 8 may notbe needed. Depending on the quantity of heavier hydrocarbons in the feedgas and the feed gas pressure, the cooled stream 33 b (FIGS. 3, 5, 6,and 7) or cooled stream 33 a (FIGS. 4 and 8) leaving heat exchanger 12may not contain any liquid (because it is above its dewpoint, or becauseit is above its cricondenbar), so that separator 13 may not bejustified. In such cases, separator 13 and expansion valve 17 may beeliminated as shown by the dashed lines. When the LNG to be processed islean or when complete vaporization of the LNG in heat exchangers 52 and53 is contemplated, separator 54 in FIGS. 3 through 8 may not bejustified. Depending on the quantity of heavier hydrocarbons in theinlet LNG and the pressure of the LNG stream leaving feed pump 51, theheated LNG stream leaving heat exchanger 53 may not contain any liquid(because it is above its dewpoint, or because it is above itscricondenbar). In such cases, separator 54 and expansion valve 59 may beeliminated as shown by the dashed lines.

In the embodiments of the present invention illustrated in FIGS. 4 and8, the expanded substantially condensed stream 32 c is formed using aportion (stream 32) of inlet gas stream 31. Depending on the feed gascomposition and other factors, some circumstances may favor using aportion of the vapor (stream 34) from separator 13 instead. In suchinstances, a portion of the separator 13 vapor forms stream 32 a asshown by the dashed lines in FIGS. 4 and 8, with the remaining portionforming the stream 34 that is fed to expansion machine 10.

In the examples shown, total condensation of stream 79 b in FIGS. 3through 5 and stream 81 b in FIGS. 6 through 8 is shown. Somecircumstances may favor subcooling these streams, while othercircumstances may favor only partial condensation. Should partialcondensation of these streams be achieved, processing of the uncondensedvapor may be necessary, using a compressor or other means to elevate thepressure of the vapor so that it can join the pumped condensed liquid.Alternatively, the uncondensed vapor could be routed to the plant fuelsystem or other such use.

Feed gas conditions, LNG conditions, plant size, available equipment, orother factors may indicate that elimination of work expansion machines10 and/or 55, or replacement with an alternate expansion device (such asan expansion valve), is feasible. Although individual stream expansionis depicted in particular expansion devices, alternative expansion meansmay be employed where appropriate.

In FIGS. 3 through 8, individual heat exchangers have been shown formost services. However, it is possible to combine two or more heatexchange services into a common heat exchanger, such as combining heatexchangers 12 and 14 in FIGS. 3 through 8 into a common heat exchanger.In some cases, circumstances may favor splitting a heat exchange serviceinto multiple exchangers. The decision as to whether to combine heatexchange services or to use more than one heat exchanger for theindicated service will depend on a number of factors including, but notlimited to, inlet gas flow rate, LNG flow rate, heat exchanger size,stream temperatures, etc. In accordance with the present invention, theuse and distribution of the methane-rich lean LNG and tower overheadstreams for process heat exchange, and the particular arrangement ofheat exchangers for heating the LNG streams and cooling the feed gasstreams, must be evaluated for each particular application, as well asthe choice of process streams for specific heat exchange services.

In the embodiments of the present invention illustrated in FIGS. 3through 8, lean LNG stream 83 a is used directly to provide cooling inheat exchanger 12 or heat exchangers 12 and 14. However, somecircumstances may favor using the lean LNG to cool an intermediate heattransfer fluid, such as propane or other suitable fluid, whereupon thecooled heat transfer fluid is then used to provide cooling in heatexchanger 12 or heat exchangers 12 and 14. This alternative means ofindirectly using the refrigeration available in lean LNG stream 83 aaccomplishes the same process objectives as the direct use of stream 83a for cooling in the FIGS. 3 through 8 embodiments of the presentinvention. The choice of how best to use the lean LNG stream forrefrigeration will depend mainly on the composition of the inlet gas,but other factors may affect the choice as well.

It will be recognized that the relative amount of feed found in eachbranch of the split LNG feed to fractionation column 62, in each branchof the split inlet gas to fractionation column 20, and in each branch ofthe split LNG feed and the split inlet gas to fractionation column 20will depend on several factors, including inlet gas composition, LNGcomposition, the amount of heat which can economically be extracted fromthe feed, and the quantity of horsepower available. More feed to the topof the column may increase recovery while increasing the duty inreboilers 61 and/or 19 and thereby increasing the high level utilityheat requirements. Increasing feed lower in the column reduces the highlevel utility heat consumption but may also reduce product recovery. Therelative locations of the mid-column feeds may vary depending on inletgas composition, LNG composition, or other factors such as the desiredrecovery level and the amount of vapor formed during heating of the LNGstreams. Moreover, two or more of the feed streams, or portions thereof,may be combined depending on the relative temperatures and quantities ofindividual streams, and the combined stream then fed to a mid-columnfeed position.

In some circumstance it may be desirable to recover refrigeration fromthe portion (stream 75 a) of LNG feed stream 71 that is fed to an uppermid-column feed point on demethanizer 62 (FIGS. 3 through 5) anddemethanizer 20 (FIGS. 6 through 8). In such cases, all of stream 71 awould be directed to heat exchanger 52 (stream 73) and the partiallyheated LNG stream (stream 73 a in FIGS. 3 through 5 and stream 73 b inFIGS. 6 through 8) would then be divided into stream 76 and stream 74(as shown by the dashed lines), whereupon stream 74 would be directed tostream 75.

In the examples given for the FIGS. 3 through 6 embodiments, recovery ofC₂ components and heavier hydrocarbon components is illustrated.However, it is believed that the FIGS. 3 through 8 embodiments are alsoadvantageous when recovery of only C₃ components and heavier hydrocarboncomponents is desired. The present invention provides improved recoveryof C₂ components and heavier hydrocarbon components or of C₃ componentsand heavier hydrocarbon components per amount of utility consumptionrequired to operate the process. An improvement in utility consumptionrequired for operating the process may appear in the form of reducedpower requirements for compression or pumping, reduced energyrequirements for tower reboilers, or a combination thereof.Alternatively, the advantages of the present invention may be realizedby accomplishing higher recovery levels for a given amount of utilityconsumption, or through some combination of higher recovery andimprovement in utility consumption.

While there have been described what are believed to be preferredembodiments of the invention, those skilled in the art will recognizethat other and further modifications may be made thereto, e.g. to adaptthe invention to various conditions, types of feed, or otherrequirements without departing from the spirit of the present inventionas defined by the following claims.

1. A process for the separation of liquefied natural gas containingmethane and heavier hydrocarbon components and a gas stream containingmethane and heavier hydrocarbon components into a volatile residue gasfraction containing a major portion of said methane and a relativelyless volatile liquid fraction containing a major portion of said heavierhydrocarbon components wherein (a) said liquefied natural gas is dividedinto at least a first liquid stream and a second liquid stream; (b) saidfirst liquid stream is expanded to lower pressure and is thereaftersupplied to a distillation column at a first upper mid-column feedposition; (c) said second liquid stream is heated sufficiently tovaporize it, thereby forming a vapor stream; (d) said vapor stream isexpanded to said lower pressure and is supplied to said distillationcolumn at a first lower mid-column feed position; (e) said gas stream isdivided into at least a first gaseous stream and a second gaseousstream; (f) said first gaseous stream is cooled to condensesubstantially all of it and is thereafter expanded to said lowerpressure whereby it is further cooled; (g) said expanded substantiallycondensed first gaseous stream is thereafter supplied to saiddistillation column at a second upper mid-column feed position; (h) saidsecond gaseous stream is expanded to said lower pressure, is cooled, andis thereafter supplied to said distillation column at a second lowermid-column feed position; (i) an overhead distillation stream iswithdrawn from an upper region of said distillation column and dividedinto at least a first portion and a second portion, whereupon said firstportion is compressed to higher pressure; (j) said compressed firstportion is cooled sufficiently to at least partially condense it andform thereby a condensed stream, with said cooling supplying at least aportion of said heating of said second liquid stream; (k) said condensedstream is divided into at least a volatile liquid stream and a refluxstream; (l) said reflux stream is further cooled, with said coolingsupplying at least a portion of said heating of said second liquidstream; (m) said further cooled reflux stream is supplied to saiddistillation column at a top column feed position; (n) said volatileliquid stream is heated sufficiently to vaporize it, with said heatingsupplying at least a portion of said cooling of one or more of saidfirst gaseous stream and said expanded second gaseous stream; (o) saidsecond portion is heated, with said heating supplying at least a portionof said cooling of one or more of said first gaseous stream and saidexpanded second gaseous stream; (p) said vaporized volatile liquidstream and said heated second portion are combined to form said volatileresidue gas fraction containing a major portion of said methane; and (q)the quantity and temperature of said reflux stream and the temperaturesof said feeds to said distillation column are effective to maintain theoverhead temperature of said distillation column at a temperaturewhereby the major portion of said heavier hydrocarbon components isrecovered in said relatively less volatile liquid fraction byfractionation in said distillation column.
 2. A process for theseparation of liquefied natural gas containing methane and heavierhydrocarbon components and a gas stream containing methane and heavierhydrocarbon components into a volatile residue gas fraction containing amajor portion of said methane and a relatively less volatile liquidfraction containing a major portion of said heavier hydrocarboncomponents wherein (a) said liquefied natural gas is divided into atleast a first liquid stream and a second liquid stream; (b) said firstliquid stream is expanded to lower pressure and is thereafter suppliedto a distillation column at a first upper mid-column feed position; (c)said second liquid stream is heated sufficiently to vaporize it, therebyforming a first vapor stream; (d) said first vapor stream is expanded tosaid lower pressure and thereafter supplied to said distillation columnat a first lower mid-column feed position; (e) said gas stream isdivided into at least a first gaseous stream and a second gaseousstream; (f) said first gaseous stream is cooled to condensesubstantially all of it and is thereafter expanded to said lowerpressure whereby it is further cooled; (g) said expanded substantiallycondensed first gaseous stream is thereafter supplied to saiddistillation column at a second upper mid-column feed position; (h) saidsecond gaseous stream is expanded to said lower pressure and isthereafter cooled sufficiently to partially condense it; (i) saidpartially condensed expanded second gaseous stream is separated therebyto provide a second vapor stream and a third liquid stream; (j) saidsecond vapor stream is further cooled and thereafter supplied to saiddistillation column at a second lower mid-column feed position; (k) saidthird liquid stream is supplied to said distillation column at a thirdlower mid-column feed position; (l) an overhead distillation stream iswithdrawn from an upper region of said distillation column and dividedinto at least a first portion and a second portion, whereupon said firstportion is compressed to higher pressure; (m) said compressed firstportion is cooled sufficiently to at least partially condense it andform thereby a condensed stream, with said cooling supplying at least aportion of said heating of said second liquid stream; (n) said condensedstream is divided into at least a volatile liquid stream and a refluxstream; (o) said reflux stream is further cooled, with said coolingsupplying at least a portion of said heating of said second liquidstream; (p) said further cooled reflux stream is supplied to saiddistillation column at a top column feed position; (q) said volatileliquid stream is heated sufficiently to vaporize it, with said heatingsupplying at least a portion of said cooling of one or more of saidfirst gaseous stream, said expanded second gaseous stream, and saidsecond vapor stream; (r) said second portion is heated, with saidheating supplying at least a portion of said cooling of one or more ofsaid first gaseous stream, said expanded second gaseous stream, and saidsecond vapor stream; (s) said vaporized volatile liquid stream and saidheated second portion are combined to form said volatile residue gasfraction containing a major portion of said methane; and (t) thequantity and temperature of said reflux stream and the temperatures ofsaid feeds to said distillation column are effective to maintain theoverhead temperature of said distillation column at a temperaturewhereby the major portion of said heavier hydrocarbon components isrecovered in said relatively less volatile liquid fraction byfractionation in said distillation column.
 3. A process for theseparation of liquefied natural gas containing methane and heavierhydrocarbon components and a gas stream containing methane and heavierhydrocarbon components into a volatile residue gas fraction containing amajor portion of said methane and a relatively less volatile liquidfraction containing a major portion of said heavier hydrocarboncomponents wherein (a) said liquefied natural gas is divided into atleast a first liquid stream and a second liquid stream; (b) said firstliquid stream is expanded to lower pressure and is thereafter suppliedto a distillation column at a first upper mid-column feed position; (c)said second liquid stream is heated sufficiently to partially vaporizeit; (d) said partially vaporized second liquid stream is separatedthereby to provide a vapor stream and a third liquid stream; (e) saidvapor stream is expanded to said lower pressure and is supplied to saiddistillation column at a first lower mid-column feed position; (f) saidgas stream is divided into at least a first gaseous stream and a secondgaseous stream; (g) said first gaseous stream is cooled to condensesubstantially all of it and is thereafter expanded to said lowerpressure whereby it is further cooled; (h) said expanded substantiallycondensed first gaseous stream is thereafter supplied to saiddistillation column at a second upper mid-column feed position; (i) saidsecond gaseous stream is expanded to said lower pressure, is cooled, andis thereafter supplied to said distillation column at a second lowermid-column feed position; (j) said third liquid stream is expanded tosaid lower pressure and thereafter supplied to said distillation columnat a third lower mid-column feed position; (k) an overhead distillationstream is withdrawn from an upper region of said distillation column anddivided into at least a first portion and a second portion, whereuponsaid first portion is compressed to higher pressure; (l) said compressedfirst portion is cooled sufficiently to at least partially condense itand form thereby a condensed stream, with said cooling supplying atleast a portion of said heating of said second liquid stream; (m) saidcondensed stream is divided into at least a volatile liquid stream and areflux stream; (n) said reflux stream is further cooled, with saidcooling supplying at least a portion of said heating of said secondliquid stream; (o) said further cooled reflux stream is supplied to saiddistillation column at a top column feed position; (p) said volatileliquid stream is heated sufficiently to vaporize it, with said heatingsupplying at least a portion of said cooling of one or more of saidfirst gaseous stream and said expanded second gaseous stream; (q) saidsecond portion is heated, with said heating supplying at least a portionof said cooling of one or more of said first gaseous stream and saidexpanded second gaseous stream; (r) said vaporized volatile liquidstream and said heated second portion are combined to form said volatileresidue gas fraction containing a major portion of said methane; and (s)the quantity and temperature of said reflux stream and the temperaturesof said feeds to said distillation column are effective to maintain theoverhead temperature of said distillation column at a temperaturewhereby the major portion of said heavier hydrocarbon components isrecovered in said relatively less volatile liquid fraction byfractionation in said distillation column.
 4. A process for theseparation of liquefied natural gas containing methane and heavierhydrocarbon components and a gas stream containing methane and heavierhydrocarbon components into a volatile residue gas fraction containing amajor portion of said methane and a relatively less volatile liquidfraction containing a major portion of said heavier hydrocarboncomponents wherein (a) said liquefied natural gas is divided into atleast a first liquid stream and a second liquid stream; (b) said firstliquid stream is expanded to lower pressure and is thereafter suppliedto a distillation column at a first upper mid-column feed position; (c)said second liquid stream is heated sufficiently to partially vaporizeit; (d) said partially vaporized second liquid stream is separatedthereby to provide a first vapor stream and a third liquid stream; (e)said first vapor stream is expanded to said lower pressure andthereafter supplied to said distillation column at a first lowermid-column feed position; (f) said gas stream is divided into at least afirst gaseous stream and a second gaseous stream; (g) said first gaseousstream is cooled to condense substantially all of it and is thereafterexpanded to said lower pressure whereby it is further cooled; (h) saidexpanded substantially condensed first gaseous stream is thereaftersupplied to said distillation column at a second upper mid-column feedposition; (i) said second gaseous stream is expanded to said lowerpressure; (j) said expanded second gaseous stream is cooled sufficientlyto partially condense it; (k) said partially condensed expanded secondgaseous stream is separated thereby to provide a second vapor stream anda fourth liquid stream; (l) said second vapor stream is further cooledand thereafter supplied to said distillation column at a second lowermid-column feed position; (m) said third liquid stream is expanded tosaid lower pressure and thereafter supplied to said distillation columnat a third lower mid-column feed position; (n) said fourth liquid streamis supplied to said distillation column at a fourth lower mid-columnfeed position; (o) an overhead distillation stream is withdrawn from anupper region of said distillation column and divided into at least afirst portion and a second portion, whereupon said first portion iscompressed to higher pressure; (p) said compressed first portion iscooled sufficiently to at least partially condense it and form thereby acondensed stream, with said cooling supplying at least a portion of saidheating of said second liquid stream; (q) said condensed stream isdivided into at least a volatile liquid stream and a reflux stream; (r)said reflux stream is further cooled, with said cooling supplying atleast a portion of said heating of said second liquid stream; (s) saidfurther cooled reflux stream is supplied to said distillation column ata top column feed position; (t) said volatile liquid stream is heatedsufficiently to vaporize it, with said heating supplying at least aportion of said cooling of one or more of said first gaseous stream,said expanded second gaseous stream, and said second vapor stream; (u)said second portion is heated, with said heating supplying at least aportion of said cooling of one or more of said first gaseous stream,said expanded second gaseous stream, and said second vapor stream; (v)said vaporized volatile liquid stream and said heated second portion arecombined to form said volatile residue gas fraction containing a majorportion of said methane; and (w) the quantity and temperature of saidreflux stream and the temperatures of said feeds to said distillationcolumn are effective to maintain the overhead temperature of saiddistillation column at a temperature whereby the major portion of saidheavier hydrocarbon components is recovered in said relatively lessvolatile liquid fraction by fractionation in said distillation column.5. The process according to claim 1 or 3 wherein (a) said second portionis compressed to higher pressure; (b) said compressed second portion isheated, with said heating supplying at least a portion of said coolingof one or more of said first gaseous stream and said expanded secondgaseous stream; and (c) said vaporized volatile liquid stream and saidheated compressed second portion are combined to form said volatileresidue gas fraction.
 6. The process according to claim 2 or 4 wherein(a) said second portion is compressed to higher pressure; (b) saidcompressed second portion is heated, with said heating supplying atleast a portion of said cooling of one or more of said first gaseousstream, said expanded second gaseous stream, and said second vaporstream; and (c) said vaporized volatile liquid stream and said heatedcompressed second portion are combined to form said volatile residue gasfraction.
 7. The process according to claim 1 or 3 wherein (a) saidsecond gaseous stream is cooled prior to said expansion; (b) said secondportion is compressed to higher pressure; (c) said volatile liquidstream is heated sufficiently to vaporize it, with said heatingsupplying at least a portion of said cooling of one or more of saidfirst gaseous stream, said second gaseous stream, and said expandedsecond gaseous stream; (d) said compressed second portion is heated,with said heating supplying at least a portion of said cooling of one ormore of said first gaseous stream, said second gaseous stream, and saidexpanded second gaseous stream; and (e) said vaporized volatile liquidstream and said heated compressed second portion are combined to formsaid volatile residue gas fraction.
 8. The process according to claim 2wherein (a) said second gaseous stream is cooled sufficiently topartially condense it; (b) said partially condensed second gaseousstream is separated thereby to provide said second vapor stream and saidthird liquid stream; (c) said second vapor stream is expanded to saidlower pressure, is cooled, and is thereafter supplied to saiddistillation column at said second lower mid-column feed position; (d)said third liquid stream is expanded to said lower pressure andthereafter supplied to said distillation column at said third lowermid-column feed position; (e) said second portion is compressed tohigher pressure; (f) said volatile liquid stream is heated sufficientlyto vaporize it, with said heating supplying at least a portion of saidcooling of one or more of said first gaseous stream, said second gaseousstream, and said expanded second vapor stream; (g) said compressedsecond portion is heated, with said heating supplying at least a portionof said cooling of one or more of said first gaseous stream, said secondgaseous stream, and said expanded second vapor stream; and (h) saidvaporized volatile liquid stream and said heated compressed secondportion are combined to form said volatile residue gas fraction.
 9. Theprocess according to claim 4 wherein (a) said second gaseous stream iscooled sufficiently to partially condense it; (b) said partiallycondensed second gaseous stream is separated thereby to provide saidsecond vapor stream and said fourth liquid stream; (c) said second vaporstream is expanded to said lower pressure, is cooled, and is thereaftersupplied to said distillation column at said second lower mid-columnfeed position; (d) said fourth liquid stream is expanded to said lowerpressure and thereafter supplied to said distillation column at saidfourth lower mid-column feed position; (e) said second portion iscompressed to higher pressure; (f) said volatile liquid stream is heatedsufficiently to vaporize it, with said heating supplying at least aportion of said cooling of one or more of said first gaseous stream,said second gaseous stream, and said expanded second vapor stream; (g)said compressed second portion is heated, with said heating supplying atleast a portion of said cooling of one or more of said first gaseousstream, said second gaseous stream, and said expanded second vaporstream; and (h) said vaporized volatile liquid stream and said heatedcompressed second portion are combined to form said volatile residue gasfraction.
 10. The process according to claim 8 wherein (a) said gasstream is cooled sufficiently to partially condense it; (b) saidpartially condensed gas stream is separated thereby to provide saidsecond vapor stream and said third liquid stream; (c) said second vaporstream is divided into at least said first gaseous stream and saidsecond gaseous stream; (d) said second gaseous stream is expanded tosaid lower pressure, is cooled, and is thereafter supplied to saiddistillation column at said second lower mid-column feed position; (e)said volatile liquid stream is heated sufficiently to vaporize it, withsaid heating supplying at least a portion of said cooling of one or moreof said gas stream, said first gaseous stream, and said expanded secondgaseous stream; and (f) said compressed second portion is heated, withsaid heating supplying at least a portion of said cooling of one or moreof said gas stream, said first gaseous stream, and said expanded secondgaseous stream.
 11. The process according to claim 9 wherein (a) saidgas stream is cooled sufficiently to partially condense it; (b) saidpartially condensed gas stream is separated thereby to provide saidsecond vapor stream and said fourth liquid stream; (c) said second vaporstream is divided into at least said first gaseous stream and saidsecond gaseous stream; (d) said second gaseous stream is expanded tosaid lower pressure, is cooled, and is thereafter supplied to saiddistillation column at said second lower mid-column feed position; (e)said volatile liquid stream is heated sufficiently to vaporize it, withsaid heating supplying at least a portion of said cooling of one or moreof said gas stream, said first gaseous stream, and said expanded secondgaseous stream; and (f) said compressed second portion is heated, withsaid heating supplying at least a portion of said cooling of one or moreof said gas stream, said first gaseous stream, and said expanded secondgaseous stream.
 12. A process for the separation of liquefied naturalgas containing methane and heavier hydrocarbon components and a gasstream containing methane and heavier hydrocarbon components into avolatile residue gas fraction containing a major portion of said methaneand a relatively less volatile liquid fraction containing a majorportion of said heavier hydrocarbon components wherein (a) saidliquefied natural gas is divided into at least a first liquid stream anda second liquid stream; (b) said first liquid stream is expanded to afirst lower pressure and is thereafter supplied to a first distillationcolumn at an upper mid-column feed position; (c) said second liquidstream is heated sufficiently to vaporize it, thereby forming a vaporstream; (d) said vapor stream is expanded to said first lower pressureand is supplied to said first distillation column at a lower mid-columnfeed position; (e) a first overhead distillation stream is withdrawnfrom an upper region of said first distillation column and compressed tohigher pressure; (f) said compressed first overhead distillation streamis cooled sufficiently to at least partially condense it and formthereby a condensed stream, with said cooling supplying at least aportion of said heating of said second liquid stream; (g) said condensedstream is divided into at least a volatile liquid stream and a refluxliquid stream; (h) said reflux liquid stream is further cooled, withsaid cooling supplying at least a portion of said heating of said secondliquid stream; (i) said further cooled reflux liquid stream is dividedinto at least a first reflux stream and a second reflux stream; (j) saidfirst reflux stream is supplied to said first distillation column at atop column feed position; (k) said gas stream is divided into at least afirst gaseous stream and a second gaseous stream; (l) said first gaseousstream is cooled to condense substantially all of it and is thereafterexpanded to a second lower pressure whereby it is further cooled; (m)said expanded substantially condensed first gaseous stream is thereaftersupplied to a second distillation column at an upper mid-column feedposition; (n) said second gaseous stream is expanded to said secondlower pressure, is cooled, and is thereafter supplied to said seconddistillation column at a lower mid-column feed position; (o) said secondreflux stream is supplied to said second distillation column at a topcolumn feed position; (p) a second overhead distillation stream iswithdrawn from an upper region of said second distillation column; (q)said volatile liquid stream is heated sufficiently to vaporize it, withsaid heating supplying at least a portion of said cooling of one or moreof said first gaseous stream and said expanded second gaseous stream;(r) said second overhead distillation stream is heated, with saidheating supplying at least a portion of said cooling of one or more ofsaid first gaseous stream and said expanded second gaseous stream; (s)said vaporized volatile liquid stream and said heated second overheaddistillation stream are combined to form said volatile residue gasfraction containing a major portion of said methane; (t) a first bottomliquid from said first distillation column and a second bottom liquidfrom said second distillation column are combined to form saidrelatively less volatile liquid fraction; and (u) the quantities andtemperatures of said first and second reflux streams and thetemperatures of said feeds to said first and second distillation columnsare effective to maintain the overhead temperatures of said first andsecond distillation columns at temperatures whereby the major portion ofsaid heavier hydrocarbon components is recovered in said relatively lessvolatile liquid fraction by fractionation in said first and seconddistillation columns.
 13. A process for the separation of liquefiednatural gas containing methane and heavier hydrocarbon components and agas stream containing methane and heavier hydrocarbon components into avolatile residue gas fraction containing a major portion of said methaneand a relatively less volatile liquid fraction containing a majorportion of said heavier hydrocarbon components wherein (a) saidliquefied natural gas is divided into at least a first liquid stream anda second liquid stream; (b) said first liquid stream is expanded to afirst lower pressure and is thereafter supplied to a first distillationcolumn at an upper mid-column feed position; (c) said second liquidstream is heated sufficiently to vaporize it, thereby forming a firstvapor stream; (d) said first vapor stream is expanded to said firstlower pressure and thereafter supplied to said first distillation columnat a lower mid-column feed position; (e) a first overhead distillationstream is withdrawn from an upper region of said first distillationcolumn and compressed to higher pressure; (f) said compressed firstoverhead distillation stream is cooled sufficiently to at leastpartially condense it and form thereby a condensed stream, with saidcooling supplying at least a portion of said heating of said secondliquid stream; (g) said condensed stream is divided into at least avolatile liquid stream and a reflux liquid stream; (h) said refluxliquid stream is further cooled, with said cooling supplying at least aportion of said heating of said second liquid stream; (i) said furthercooled reflux liquid stream is divided into at least a first refluxstream and a second reflux stream; (j) said first reflux stream issupplied to said first distillation column at a top column feedposition; (k) said gas stream is divided into at least a first gaseousstream and a second gaseous stream; (l) said first gaseous stream iscooled to condense substantially all of it and is thereafter expanded toa second lower pressure whereby it is further cooled; (m) said expandedsubstantially condensed first gaseous stream is thereafter supplied to asecond distillation column at an upper mid-column feed position; (n)said second gaseous stream is expanded to said second lower pressure andis thereafter cooled sufficiently to partially condense it; (o) saidpartially condensed expanded second gaseous stream is separated therebyto provide a second vapor stream and a third liquid stream; (p) saidsecond vapor stream is further cooled and thereafter supplied to saidsecond distillation column at a first lower mid-column feed position;(q) said third liquid stream is supplied to said second distillationcolumn at a second lower mid-column feed position; (r) said secondreflux stream is supplied to said second distillation column at a topcolumn feed position; (s) a second overhead distillation stream iswithdrawn from an upper region of said second distillation column; (t)said volatile liquid stream is heated sufficiently to vaporize it, withsaid heating supplying at least a portion of said cooling of one or moreof said first gaseous stream, said expanded second gaseous stream, andsaid second vapor stream; (u) said second overhead distillation streamis heated, with said heating supplying at least a portion of saidcooling of one or more of said first gaseous stream, said expandedsecond gaseous stream, and said second vapor stream; (v) said vaporizedvolatile liquid stream and said heated second overhead distillationstream are combined to form said volatile residue gas fractioncontaining a major portion of said methane; (w) a first bottom liquidfrom said first distillation column and a second bottom liquid from saidsecond distillation column are combined to form said relatively lessvolatile liquid fraction; and (x) the quantities and temperatures ofsaid first and second reflux streams and the temperatures of said feedsto said first and second distillation columns are effective to maintainthe overhead temperatures of said first and second distillation columnsat temperatures whereby the major portion of said heavier hydrocarboncomponents is recovered in said relatively less volatile liquid fractionby fractionation in said first and second distillation columns.
 14. Aprocess for the separation of liquefied natural gas containing methaneand heavier hydrocarbon components and a gas stream containing methaneand heavier hydrocarbon components into a volatile residue gas fractioncontaining a major portion of said methane and a relatively lessvolatile liquid fraction containing a major portion of said heavierhydrocarbon components wherein (a) said liquefied natural gas is dividedinto at least a first liquid stream and a second liquid stream; (b) saidfirst liquid stream is expanded to a first lower pressure and isthereafter supplied to a first distillation column at an uppermid-column feed position; (c) said second liquid stream is heatedsufficiently to partially vaporize it; (d) said partially vaporizedsecond liquid stream is separated thereby to provide a vapor stream anda third liquid stream; (e) said vapor stream is expanded to said firstlower pressure and is supplied to said first distillation column at afirst lower mid-column feed position; (f) said third liquid stream isexpanded to said first lower pressure and thereafter supplied to saidfirst distillation column at a second lower mid-column feed position;(g) a first overhead distillation stream is withdrawn from an upperregion of said first distillation column and compressed to higherpressure; (h) said compressed first overhead distillation stream iscooled sufficiently to at least partially condense it and form thereby acondensed stream, with said cooling supplying at least a portion of saidheating of said second liquid stream; (i) said condensed stream isdivided into at least a volatile liquid stream and a reflux liquidstream; (j) said reflux liquid stream is further cooled, with saidcooling supplying at least a portion of said heating of said secondliquid stream; (k) said further cooled reflux liquid stream is dividedinto at least a first reflux stream and a second reflux stream; (l) saidfirst reflux stream is supplied to said first distillation column at atop column feed position; (m) said gas stream is divided into at least afirst gaseous stream and a second gaseous stream; (n) said first gaseousstream is cooled to condense substantially all of it and is thereafterexpanded to a second lower pressure whereby it is further cooled; (o)said expanded substantially condensed first gaseous stream is thereaftersupplied to a second distillation column at an upper mid-column feedposition; (p) said second gaseous stream is expanded to said secondlower pressure, is cooled, and is thereafter supplied to said seconddistillation column at a lower mid-column feed position; (q) said secondreflux stream is supplied to said second distillation column at a topcolumn feed position; (r) a second overhead distillation stream iswithdrawn from an upper region of said second distillation column; (s)said volatile liquid stream is heated sufficiently to vaporize it, withsaid heating supplying at least a portion of said cooling of one or moreof said first gaseous stream and said expanded second gaseous stream;(t) said second overhead distillation stream is heated, with saidheating supplying at least a portion of said cooling of one or more ofsaid first gaseous stream and said expanded second gaseous stream; (u)said vaporized volatile liquid stream and said heated second overheaddistillation stream are combined to form said volatile residue gasfraction containing a major portion of said methane; (v) a first bottomliquid from said first distillation column and a second bottom liquidfrom said second distillation column are combined to form saidrelatively less volatile liquid fraction; and (w) the quantities andtemperatures of said first and second reflux streams and thetemperatures of said feeds to said first and second distillation columnsare effective to maintain the overhead temperatures of said first andsecond distillation columns at temperatures whereby the major portion ofsaid heavier hydrocarbon components is recovered in said relatively lessvolatile liquid fraction by fractionation in said first and seconddistillation columns.
 15. A process for the separation of liquefiednatural gas containing methane and heavier hydrocarbon components and agas stream containing methane and heavier hydrocarbon components into avolatile residue gas fraction containing a major portion of said methaneand a relatively less volatile liquid fraction containing a majorportion of said heavier hydrocarbon components wherein (a) saidliquefied natural gas is divided into at least a first liquid stream anda second liquid stream; (b) said first liquid stream is expanded to afirst lower pressure and is thereafter supplied to a first distillationcolumn at an upper mid-column feed position; (c) said second liquidstream is heated sufficiently to partially vaporize it; (d) saidpartially vaporized second liquid stream is separated thereby to providea first vapor stream and a third liquid stream; (e) said first vaporstream is expanded to said first lower pressure and thereafter suppliedto said first distillation column at a first lower mid-column feedposition; (f) said third liquid stream is expanded to said first lowerpressure and thereafter supplied to said first distillation column at asecond lower mid-column feed position; (g) a first overhead distillationstream is withdrawn from an upper region of said first distillationcolumn and compressed to higher pressure; (h) said compressed firstoverhead distillation stream is cooled sufficiently to at leastpartially condense it and form thereby a condensed stream, with saidcooling supplying at least a portion of said heating of said secondliquid stream; (i) said condensed stream is divided into at least avolatile liquid stream and a reflux liquid stream; (j) said refluxliquid stream is further cooled, with said cooling supplying at least aportion of said heating of said second liquid stream; (k) said furthercooled reflux liquid stream is divided into at least a first refluxstream and a second reflux stream; (l) said first reflux stream issupplied to said first distillation column at a top column feedposition; (m) said gas stream is divided into at least a first gaseousstream and a second gaseous stream; (n) said first gaseous stream iscooled to condense substantially all of it and is thereafter expanded toa second lower pressure whereby it is further cooled; (o) said expandedsubstantially condensed first gaseous stream is thereafter supplied to asecond distillation column at an upper mid-column feed position; (p)said second gaseous stream is expanded to said second lower pressure andis thereafter cooled sufficiently to partially condense it; (q) saidpartially condensed expanded second gaseous stream is separated therebyto provide a second vapor stream and a fourth liquid stream; (r) saidsecond vapor stream is further cooled and thereafter supplied to saidsecond distillation column at a first lower mid-column feed position;(s) said fourth liquid stream is supplied to said second distillationcolumn at a second lower mid-column feed position; (t) said secondreflux stream is supplied to said second distillation column at a topcolumn feed position; (u) a second overhead distillation stream iswithdrawn from an upper region of said second distillation column; (v)said volatile liquid stream is heated sufficiently to vaporize it, withsaid heating supplying at least a portion of said cooling of one or moreof said first gaseous stream, said expanded second gaseous stream, andsaid second vapor stream; (w) said second overhead distillation streamis heated, with said heating supplying at least a portion of saidcooling of one or more of said first gaseous stream, said expandedsecond gaseous stream, and said second vapor stream; (x) said vaporizedvolatile liquid stream and said heated second overhead distillationstream are combined to form said volatile residue gas fractioncontaining a major portion of said methane; (y) a first bottom liquidfrom said first distillation column and a second bottom liquid from saidsecond distillation column are combined to form said relatively lessvolatile liquid fraction; and (z) the quantities and temperatures ofsaid first and second reflux streams and the temperatures of said feedsto said first and second distillation columns are effective to maintainthe overhead temperatures of said first and second distillation columnsat temperatures whereby the major portion of said heavier hydrocarboncomponents is recovered in said relatively less volatile liquid fractionby fractionation in said first and second distillation columns.
 16. Theprocess according to claim 12 or 14 wherein (a) said second overheaddistillation stream is compressed to higher pressure; (b) saidcompressed second overhead distillation stream is heated, with saidheating supplying at least a portion of said cooling of one or more ofsaid first gaseous stream and said expanded second gaseous stream; and(c) said vaporized volatile liquid stream and said heated compressedsecond overhead distillation stream are combined to form said volatileresidue gas fraction.
 17. The process according to claim 13 or 15wherein (a) said second overhead distillation stream is compressed tohigher pressure; (b) said compressed second overhead distillation streamis heated, with said heating supplying at least a portion of saidcooling of one or more of said first gaseous stream, said expandedsecond gaseous stream, and said second vapor stream; and (c) saidvaporized volatile liquid stream and said heated compressed secondoverhead distillation stream are combined to form said volatile residuegas fraction.
 18. The process according to claim 12 or 14 wherein (a)said second gaseous stream is cooled prior to said expansion; (b) saidsecond overhead distillation stream is compressed to higher pressure;(c) said volatile liquid stream is heated sufficiently to vaporize it,with said heating supplying at least a portion of said cooling of one ormore of said first gaseous stream, said second gaseous stream, and saidexpanded second gaseous stream; (d) said compressed second overheaddistillation stream is heated, with said heating supplying at least aportion of said cooling of one or more of said first gaseous stream,said second gaseous stream, and said expanded second gaseous stream; and(e) said vaporized volatile liquid stream and said heated compressedsecond overhead distillation stream are combined to form said volatileresidue gas fraction.
 19. The process according to claim 13 wherein (a)said second gaseous stream is cooled sufficiently to partially condenseit; (b) said partially condensed second gaseous stream is separatedthereby to provide said second vapor stream and said third liquidstream; (c) said second vapor stream is expanded to said second lowerpressure, is cooled, and is thereafter supplied to said seconddistillation column at said first lower mid-column feed position; (d)said third liquid stream is expanded to said second lower pressure andthereafter supplied to said second distillation column at said secondlower mid-column feed position; (e) said second overhead distillationstream is compressed to higher pressure; (f) said volatile liquid streamis heated sufficiently to vaporize it, with said heating supplying atleast a portion of said cooling of one or more of said first gaseousstream, said second gaseous stream, and said expanded second vaporstream; (g) said compressed second overhead distillation stream isheated, with said heating supplying at least a portion of said coolingof one or more of said first gaseous stream, said second gaseous stream,and said expanded second vapor stream; and (h) said vaporized volatileliquid stream and said heated compressed second overhead distillationstream are combined to form said volatile residue gas fraction.
 20. Theprocess according to claim 15 wherein (a) said second gaseous stream iscooled sufficiently to partially condense it; (b) said partiallycondensed second gaseous stream is separated thereby to provide saidsecond vapor stream and said fourth liquid stream; (c) said second vaporstream is expanded to said second lower pressure, is cooled, and isthereafter supplied to said second distillation column at said firstlower mid-column feed position; (d) said fourth liquid stream isexpanded to said second lower pressure and thereafter supplied to saidsecond distillation column at said second lower mid-column feedposition; (e) said second overhead distillation stream is compressed tohigher pressure; (f) said volatile liquid stream is heated sufficientlyto vaporize it, with said heating supplying at least a portion of saidcooling of one or more of said first gaseous stream, said second gaseousstream, and said expanded second vapor stream; (g) said compressedsecond overhead distillation stream is heated, with said heatingsupplying at least a portion of said cooling of one or more of saidfirst gaseous stream, said second gaseous stream, and said expandedsecond vapor stream; and (h) said vaporized volatile liquid stream andsaid heated compressed second overhead distillation stream are combinedto form said volatile residue gas fraction.
 21. The process according toclaim 19 wherein (a) said gas stream is cooled sufficiently to partiallycondense it; (b) said partially condensed gas stream is separatedthereby to provide said second vapor stream and said third liquidstream; (c) said second vapor stream is divided into at least said firstgaseous stream and said second gaseous stream; (d) said second gaseousstream is expanded to said second lower pressure, is cooled, and isthereafter supplied to said second distillation column at said firstlower mid-column feed position; (e) said volatile liquid stream isheated sufficiently to vaporize it, with said heating supplying at leasta portion of said cooling of one or more of said gas stream, said firstgaseous stream, and said expanded second gaseous stream; and (f) saidcompressed second overhead distillation stream is heated, with saidheating supplying at least a portion of said cooling of one or more ofsaid gas stream, said first gaseous stream, and said expanded secondgaseous stream.
 22. The process according to claim 20 wherein (a) saidgas stream is cooled sufficiently to partially condense it; (b) saidpartially condensed gas stream is separated thereby to provide saidsecond vapor stream and said fourth liquid stream; (c) said second vaporstream is divided into at least said first gaseous stream and saidsecond gaseous stream; (d) said second gaseous stream is expanded tosaid second lower pressure, is cooled, and is thereafter supplied tosaid second distillation column at said first lower mid-column feedposition; (e) said volatile liquid stream is heated sufficiently tovaporize it, with said heating supplying at least a portion of saidcooling of one or more of said gas stream, said first gaseous stream,and said expanded second gaseous stream; and (f) said compressed secondoverhead distillation stream is heated, with said heating supplying atleast a portion of said cooling of one or more of said gas stream, saidfirst gaseous stream, and said expanded second gaseous stream.
 23. Theprocess according to claim 1, 2, 3, 4, 8, 9, 10, or 11 wherein (a) saidliquefied natural gas is heated and thereafter divided into at leastsaid first liquid stream and said second liquid stream; and (b) saidcooling of said compressed first portion and said reflux stream supplyat least a portion of said heating of said liquefied natural gas. 24.The process according to claim 5 wherein (a) said liquefied natural gasis heated and thereafter divided into at least said first liquid streamand said second liquid stream; and (b) said cooling of said compressedfirst portion and said reflux stream supply at least a portion of saidheating of said liquefied natural gas.
 25. The process according toclaim 6 wherein (a) said liquefied natural gas is heated and thereafterdivided into at least said first liquid stream and said second liquidstream; and (b) said cooling of said compressed first portion and saidreflux stream supply at least a portion of said heating of saidliquefied natural gas.
 26. The process according to claim 7 wherein (a)said liquefied natural gas is heated and thereafter divided into atleast said first liquid stream and said second liquid stream; and (b)said cooling of said compressed first portion and said reflux streamsupply at least a portion of said heating of said liquefied natural gas.27. The process according to claim 12, 13, 14, 15, 19, 20, 21, or 22wherein (a) said liquefied natural gas is heated and thereafter dividedinto at least said first liquid stream and said second liquid stream;and (b) said cooling of said compressed first overhead distillationstream and said reflux liquid stream supply at least a portion of saidheating of said liquefied natural gas.
 28. The process according toclaim 16 wherein (a) said liquefied natural gas is heated and thereafterdivided into at least said first liquid stream and said second liquidstream; and (b) said cooling of said compressed first overheaddistillation stream and said reflux liquid stream supply at least aportion of said heating of said liquefied natural gas.
 29. The processaccording to claim 17 wherein (a) said liquefied natural gas is heatedand thereafter divided into at least said first liquid stream and saidsecond liquid stream; and (b) said cooling of said compressed firstoverhead distillation stream and said reflux liquid stream supply atleast a portion of said heating of said liquefied natural gas.
 30. Theprocess according to claim 18 wherein (a) said liquefied natural gas isheated and thereafter divided into at least said first liquid stream andsaid second liquid stream; and (b) said cooling of said compressed firstoverhead distillation stream and said reflux liquid stream supply atleast a portion of said heating of said liquefied natural gas.
 31. Theprocess according to claim 1, 2, 3, 4, 8, 9, 10, 11, 12, 13, 14, 15, 19,20, 21, or 22 wherein said volatile residue gas fraction contains amajor portion of said methane and C₂ components.
 32. The processaccording to claim 5 wherein said volatile residue gas fraction containsa major portion of said methane and C₂ components.
 33. The processaccording to claim 6 wherein said volatile residue gas fraction containsa major portion of said methane and C₂ components.
 34. The processaccording to claim 7 wherein said volatile residue gas fraction containsa major portion of said methane and C₂ components.
 35. The processaccording to claim 16 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components
 36. Theprocess according to claim 17 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 37. Theprocess according to claim 18 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 38. Theprocess according to claim 23 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 39. Theprocess according to claim 24 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 40. Theprocess according to claim 25 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 41. Theprocess according to claim 26 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 42. Theprocess according to claim 27 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 43. Theprocess according to claim 28 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 44. Theprocess according to claim 29 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.
 45. Theprocess according to claim 30 wherein said volatile residue gas fractioncontains a major portion of said methane and C₂ components.